Integrated biorefinery for production of liquid fuels

ABSTRACT

A system including a mixing apparatus configured to produce a reformer feedstock and comprising one or more cylindrical vessel having a conical bottom section, an inlet for superheated steam within the conical bottom section and an inlet for at least one carbonaceous material at or near the top of the cylindrical vessel, wherein the one or more cylindrical vessel is a pressure vessel configured for operation at a pressure in the range of from about 5 psig (34.5 kPa) to about 50 psig (344.7 kPa); a reformer configured to produce, from the reformer feedstock, a reformer product comprising synthesis gas, and also producing a hot flue gas; a synthesis gas conversion apparatus configured to catalytically convert at least a portion of the synthesis gas in the reformer product into synthesis gas conversion product, and to separate, from the synthesis gas conversion product, a spent catalyst stream and a tailgas.

CROSS-REFERENCE TO RELATED APPLICATIONS

This application is a continuation of U.S. patent application Ser. No.13/246,451, filed Sep. 27, 2011, which is itself a divisionalapplication which claims the benefit under 35 U.S.C. § 121 of U.S.patent application Ser. No. 12/976,763, (now U.S. Pat. No. 8,168,686),filed Dec. 22, 2010, the disclosure of each of which is herebyincorporated herein by reference in its entirety for all purposes.

STATEMENT REGARDING FEDERALLY SPONSORED RESEARCH OR DEVELOPMENT

Not Applicable.

BACKGROUND

Field of the Invention

This disclosure relates generally to a biorefinery and method for theconversion of carbonaceous feedstock into synthesis gas conversionproducts. More specifically, this disclosure relates to a biorefineryand method for the conversion of carbonaceous feedstock to liquidhydrocarbons via Fischer-Tropsch. Still more specifically, thisdisclosure relates to a biorefinery and method for the conversion ofcarbonaceous material to Fischer-Tropsch products wherein at least onebyproduct of Fischer-Tropsch conversion is utilized to produceadditional synthesis gas for Fischer-Tropsch synthesis.

Background of the Invention

Processes for the production of synthesis gas from carbonaceousmaterials utilize gasification of a feedstock comprising thecarbonaceous materials in a so-called ‘reformer’ to produce a streamcomprising synthesis gas (i.e. hydrogen and carbon monoxide; also knownas ‘syngas’). The product synthesis gas generally also comprises amountsof carbon dioxide and methane and may also comprise minor amounts ofother components. Generation of synthesis gas is disclosed in numerouspatents.

Synthesis gas produced via gasification of carbonaceous materials can beconverted into other compounds in a so-called Fischer-Tropsch reaction.Fischer-Tropsch (FT) synthesis can be used to catalytically producesynthetic liquid fuels, alcohols or other oxidized compounds. FTsynthesis occurs by the metal catalysis of an exothermic reaction ofsynthesis gas. Fischer-Tropsch (FT) technology can thus be utilized toconvert synthesis gas to valuable products. Hydrocarbon liquid productsof various Fischer-Tropsch processes are generally refined to produce arange of synthetic fuels, lubricants and waxes. Often, theFischer-Tropsch process is performed in a slurry bubble column reactor(SBCR). The technology of converting synthesis gas originating fromnatural gas into valuable primarily liquid hydrocarbon products isreferred to as Gas To Liquids (GTL) technology. When coal is the rawmaterial for the syngas, the technology is commonly referred to asCoal-To-Liquids (CTL). Fischer-Tropsch technology is one of severalconversion techniques included in the broader GTL/CTL technology.Desirably, the synthesis gas for subsequent production of valuableproducts via Fischer-Tropsch is produced from ‘green’ materials. Forexample, an environmentally-friendly system for the production ofsynthesis gas, which may subsequently be utilized to produceFischer-Tropsch products, would desirably allow for the production ofsynthesis gas from carbonaceous materials, such as biomass, which maygenerally be considered waste materials

The catalyst used in the Fischer-Tropsch reactor and, to some extent,the temperatures and pressures used will determine what products can beobtained. Some Fischer-Tropsch processes are directed to the productionof liquid hydrocarbons. Such processes generally utilize iron-,ruthenium, or cobalt-based catalysts. Iron-based catalysts are generallyoperated with a synthesis gas having a molar ratio of hydrogen to carbonmonoxide in the range of from about 0.7 to about 2.0. Cobalt-basedcatalysts are generally operated with a synthesis gas having a moleratio of hydrogen to carbon monoxide in the range of from about 1.8 toabout 2.2. For example, carbon monoxide and hydrogen can be converted toalkanes over a cobalt-thoria catalyst. U.S. Pat. No. 4,609,679 teachesthe use of ruthenium combined with tantalum, niobium, vanadium ormixtures thereof to selectively catalyze for the production of methane.As mentioned hereinabove, other Fischer-Tropsch processes are directedtoward the production of alcohols.

Accordingly, there is a need in the art for systems and methods for theproduction of synthesis gas conversion products from carbonaceousmaterials. Such systems and methods should preferably enable theenvironmentally-friendly production of synthesis gas conversion product,for example, by allowing the production of synthesis gas fromsustainable and renewable feedstocks such as biomass, facilitatingsequestration of carbon dioxide, and/or reducing the amount of wastematerial produced.

SUMMARY

Herein disclosed are a system and method of producing synthesis gasconversion product. Herein disclosed is a method of producing liquidhydrocarbons, the method comprising: reforming a carbonaceous feedstockthat is solid, liquid, or both to produce a first synthesis gascomprising hydrogen and carbon monoxide; subjecting at least a portionof the first synthesis gas to Fischer-Tropsch conversion whereby atleast a portion of the hydrogen and carbon monoxide in the firstsynthesis gas is catalytically converted into product comprising liquidhydrocarbons; separating from the product a Fischer-Tropsch tailgascomprising at least one component selected from carbon monoxide,hydrogen, methane and carbon dioxide; and combusting at least a portionof the Fischer-Tropsch tailgas to provide at least a portion of the heatfor the reforming of additional carbonaceous feedstock. The method mayfurther comprise forming additional carbonaceous feedstock by combining,with superheated steam, at least one carbonaceous material and a spentcatalyst stream comprising Fischer-Tropsch liquid hydrocarbons andcatalyst that has been at least partially deactivated, attrited, or bothduring Fischer-Tropsch conversion.

Also disclosed herein is a method of producing liquid hydrocarbons, themethod comprising: reforming a carbonaceous feedstock to produce a firstsynthesis gas comprising hydrogen and carbon monoxide; subjecting atleast a portion of the first synthesis gas to Fischer-Tropsch conversionwhereby at least a portion of the hydrogen and carbon monoxide in thefirst synthesis gas is catalytically converted into product comprisingliquid hydrocarbons; removing from the Fischer-Tropsch conversionreactor a catalyst wax mixture comprising Fischer-Tropsch liquidhydrocarbons and catalyst removed from the reactor; and combining atleast a portion of the catalyst wax mixture with at least onecarbonaceous material and superheated steam; and reforming the combinedmaterial to produce additional synthesis gas. In embodiments, the methodfurther comprises combusting at least a portion of a Fischer-Tropschtailgas produced during Fischer-Tropsch conversion and comprising atleast one component selected from the group consisting of carbonmonoxide, hydrogen, carbon dioxide and methane to provide heat forreforming of additional carbonaceous feedstock.

Also disclosed herein is a system for the production of conversionproducts from synthesis gas, the system comprising: a mixing apparatusconfigured for mixing steam with at least one carbonaceous material toproduce a reformer feedstock; a reformer configured to produce, from thereformer feedstock, a reformer product comprising synthesis gascomprising hydrogen and carbon monoxide from the reformer feedstock; asynthesis gas conversion apparatus configured to catalytically convertat least a portion of the synthesis gas in the reformer product intosynthesis gas conversion product and to separate from the synthesis gasconversion product a tailgas comprising at least one gas selected fromthe group consisting of carbon monoxide, carbon dioxide, hydrogen andmethane; and one or more recycle lines fluidly connecting the synthesisgas conversion apparatus with the mixing apparatus, the reformer, orboth. In embodiments, the synthesis gas conversion apparatus isconfigured to catalytically convert at least a portion of the synthesisgas in the reformer product into synthesis gas conversion product viacontact of the at least a portion of the synthesis gas with aniron-based Fischer-Tropsch catalyst.

In embodiments, the system comprises a recycle line fluidly connectingthe synthesis gas conversion apparatus with at least one burner of thereformer, whereby at least a portion of the tailgas can be combusted toprovide heat. In embodiments, the at least one carbonaceous materialcomprises biomass.

In embodiments, the mixing apparatus is a pressure vessel operable at apressure of about 5 psig (34.5 kPa) to 45 psig (310.3 kPa). Inembodiments, the mixing apparatus comprises one or more cylindricalvessels having a conical bottom section, an inlet for superheated steamwithin the conical bottom section and an inlet for the at least onecarbonaceous material at or near the top of the cylindrical vessel.

In embodiments the metallurgy of the reformer allows operation at areformer temperature greater than or equal to about 1700° F. (926° C.)and a reformer pressure greater than or equal to about 5 psig (34.5kPa). In embodiments of a system as disclosed herein, the reformercomprises: a cylindrical vessel containing a plurality ofvertically-oriented coiled tubes fluidly connected with the mixingapparatus such that reformer feedstock may be introduced thereto; atleast one burner configured to combust fuel to provide heat for thereforming and produce a flue gas; at least one outlet for reformerproduct; and at least one outlet for the flue gas.

In embodiments, each of the plurality of vertically-oriented coiledtubes has a vertical height in the range of from about 40 feet (12.2 m)to about 100 feet (30.5 m) and a coil length at least 4 times thevertical height. In embodiments, the total coil length is in the rangeof from about 4 to about 25 times the vertical height. In embodiments,the total coil length is in the range of from about 4 to about 12 timesthe vertical height. In embodiments, at least a portion of the pluralityof vertically-oriented coiled tubes have an inside diameter (ID) in therange of from about 2 inches (5.1 cm) to about 4 inches (10.2 cm). Inembodiments, the metallurgy of the coiled tubes allows operation at areformer pressure greater than or equal to about 45 psig (310.3 kPa).

In embodiments, the at least one burner is positioned substantially at,near, or below the bottom of the cylindrical vessel. In embodiments, thereformer comprises from about 1 to about 10 burners. In embodiments,outlets of each of the coiled tubes are manifolded into an outlet forthe reformer product, wherein the manifold is positioned at, near, orbelow the bottom of the cylindrical vessel. In embodiments according tothe disclosure, the at least one outlet for flue gas is positioned atthe top of the cylindrical vessel. In embodiments, the system furthercomprises a steam superheater configured to produce superheated steamvia heat transfer from the reformer flue gas.

In embodiments, the system further comprises feed preparation apparatusconfigured to comminute the at least one carbonaceous material, to drythe at least one carbonaceous material, or both. In embodiments, thefeed preparation apparatus comprises at least one grinder and at leastone separator configured to provide a carbonaceous material having anaverage particle diameter of less than about 3/16^(th) of an inch (0.47cm).

The foregoing has outlined rather broadly the features and technicaladvantages of the invention in order that the detailed description ofthe invention that follows may be better understood. Additional featuresand advantages of the invention will be described hereinafter that formthe subject of the claims of the invention. It should be appreciated bythose skilled in the art that the conception and the specificembodiments disclosed may be readily utilized as a basis for modifyingor designing other structures for carrying out the same purposes of theinvention. It should also be realized by those skilled in the art thatsuch equivalent constructions do not depart from the spirit and scope ofthe invention as set forth in the appended claims.

BRIEF DESCRIPTION OF THE DRAWINGS

For a more detailed description of embodiments of the present invention,reference will now be made to the accompanying drawings, wherein:

FIG. 1 is a schematic of a biorefinery suitable for carrying out theproduction of synthesis gas conversion products according to anembodiment of this disclosure;

FIG. 2 is a schematic of suitable mixing apparatus, biomass reformer,and steam generation apparatus for use in the biorefinery of FIG. 1according to an embodiment of this disclosure;

FIG. 3 is a schematic of suitable mixing apparatus, biomass reformer,and steam generation apparatus for use in the biorefinery of FIG. 1according to another embodiment of this disclosure;

FIG. 4 is a schematic of a synthesis gas cleanup and/or conditioningapparatus suitable for use in the biorefinery of FIG. 1 according to anembodiment of this disclosure;

FIG. 5 is a schematic of a feedstock handling and/or drying apparatussuitable for use in the biorefinery of FIG. 1 according to an embodimentof this disclosure;

FIG. 6 is a flow diagram of a method of producing synthesis gasconversion product(s) according to an embodiment of this disclosure;

FIG. 7 is a flow diagram of a method of producing synthesis gasaccording to an embodiment of this disclosure; and

FIG. 8 is a flow diagram of a method for converting synthesis gas toproduct according to an embodiment of this disclosure.

NOTATION AND NOMENCLATURE

Certain terms are used throughout the following description and claimsto refer to particular system components. This document does not intendto distinguish between components that differ in name but not function.

As used herein, the term ‘carbonaceous feedstock’ includes not onlyorganic matter that is part of the stable carbon cycle, but alsofossilized organic matter such as coal, petroleum, and natural gas, andproducts, derivatives and byproducts thereof, such as plastics,petroleum coke and the like.

As used herein, the terms ‘hot’, ‘warm’, ‘cool’ and ‘cold’ are utilizedto refer to the relative condition of various streams. That is, a ‘hot’stream is at a higher temperature than a ‘warm’ stream, a ‘warm’ streamis likewise at a higher temperature than a ‘cool’ stream and a ‘cool’stream is likewise at a higher temperature than a ‘cold’ stream. Such astream may not normally be considered as such. That is a ‘cool’ streammay have a temperature that is actually high enough to be considered hotor warm in conventional, non-relative usage.

As used herein the term ‘dry’ as applied to a carbonaceous feed materialis used to indicate that the feed material has a moisture contentsuitable for reforming, e.g. less than about 20 weight percent, and notto imply the complete absence of moisture.

DETAILED DESCRIPTION I. Overview

Herein disclosed are a biorefinery and a method for producing synthesisgas conversion products such as, but not limited to, Fischer-Tropschhydrocarbons. The disclosed biorefinery and method enable the use ofrenewable and sustainable carbonaceous materials, such as biomass, forthe production of synthesis gas, the sequestration of carbon dioxide inmultiple ways and locations, a reduction in the amount of waste fordisposal (e.g. Fischer-Tropsch wax associated with spent catalyst), anda reduction in the amount of ‘waste’ tailgas. Accordingly, the disclosedbiorefinery and process for producing synthesis gas conversion productstherewith represent clean technologies. Such a biorefinery issignificantly more environmentally-friendly than conventionalbiorefineries that produce synthesis gas for subsequent conversion fromother sources, such as from natural gas.

II. Biorefinery

FIG. 1 is a schematic of a biorefinery 100 according to this disclosure.Biorefinery 100 comprises mixing apparatus 300, reforming apparatus 400,steam generation apparatus 500 and synthesis gas conversion apparatus700. As discussed further hereinbelow, biorefinery 100 can furthercomprise feed handling and/or drying apparatus 200 and synthesis gascleanup and/or conditioning apparatus 600. Each of the basic apparatuswill be described in more detail hereinbelow.

Reforming Apparatus 400.

Biorefinery 100 comprises reforming apparatus 400 (also at timesreferred to herein as ‘reformer 400’). Description of reformingapparatus 400 will now be made with reference to FIG. 2, which is aschematic of a portion 100A of a biorefinery comprising mixing apparatus300A, reformer 400A and steam generation apparatus 500A, according to anembodiment of this disclosure, and FIG. 3, which is a schematic of aportion 100B of a biorefinery comprising mixing apparatus 300B, reformer400B and steam generation apparatus 500B, according to anotherembodiment of this disclosure.

Reformer 400A is a high temperature, high efficiency reformer. Inembodiments, reformer 400 is a biomass reformer. Reformer 400A comprisesa plurality of coiled tubes 410A, 410B surrounded by enclosure,cylindrical vessel or firebox 407. In embodiments, biomass reformer 400Ais a cylindrical vessel. In embodiments, the cylindrical vessel 407 hasa height H1 in the range of from about 40 feet (12.2 m) to about 100feet (30.5 m), from about 50 feet (15.2 m) to about 100 feet (30.5 m),or from about 60 feet (18.3 m) to about 100 feet (30.5 m). Inembodiments, coiled tubes 410 have an inside diameter (ID) of at leastor about 2 inches (5.1 cm), at least or about 3 inches (7.6 cm), or atleast or about 4 inches (10.2 cm). Coiled tubes 410 may be configured ascylindrical helices and may be oriented vertically within cylindricalvessel 407. In embodiments, each of the coiled tubes 410 has a totallength or coil length that is at least 4, 5, 10, 15, 20 or 25 times thevertical height of the coiled tubes. In embodiments, each of the coiledtubes 410 has a total length in the range of from about 200 feet (61 m)to about 900 feet (274 m), from about 300 feet (91.4 m) to about 700feet (213.4 m), or from about 350 feet (106.7 m) to about 650 feet(198.1 m).

In embodiments, the metallurgy of the coiled tubes is upgraded such thatthe tubes are operable at the high temperatures of operation of a hightemperature reformer. A ‘high’ temperature reformer is operable at atemperature of at least 1093° C. (2000° F.). In embodiments, the coiledtubes are operable at temperatures up to 926° C. (1700° F.), 982° C.(1800° F.), 1038° C. (1900° F.), 1093° C. (2000° F.), 1149° C. (2100°F.) and a pressure of at least 2 psig (13.8 kPa), 5 psig (34.5 kPa), atleast 20 psig (137.9 kPa), greater than or about 40 psig (275.8 kPa) orabout 45 psig (310.3 kPa) or about 50 psig (344.7 kPa). In embodiments,the coiled tubes are fabricated from stainless steel or other high alloysteel, such as 310 stainless steel. In embodiments, the coiled tubes arefabricated from austenitic nickel-chromium-based superalloys or otherhigh temperature alloys that are resistant to hydrogen attack andsuitable for production of coiled helices, such as INCONEL™. Inembodiments, the coiled tubes are fabricated from INCONEL™ 800 HT. Inembodiments, the coiled tubes are designed to provide at least 100,000hours of operation.

As shown in FIG. 3, a distributor or flow divider 412 can be positionedexternal or internal to firebox 407 for distributing a reformerfeedstock comprising a mixture of cooled steam and dry carbonaceousmaterial to the plurality of coiled tubes 410. In embodiments,distributor 412 is positioned external to vessel 407. In embodiments,distributor 412 is configured to provide substantially equal amounts ofthe reformer feed mixture to each of the coiled tubes 410.

Distributor 412 distributes reformer feed mixture to each of theplurality of coiled tubes 410 (410A and 410B indicated in the embodimentof FIG. 3) via one or more reformer feed inlet lines 350 (350A and 350Bdepicted in the embodiment of FIG. 3). In embodiments, mixing apparatus300 (300A in FIG. 2; 300B in FIG. 3), further discussed hereinbelow,comprises a plurality of feed mixers 310 (mixers 310A and 310B depictedin FIG. 2; mixer 310C depicted in FIG. 3), the output of each of whichis fed via one or more reformer feed inlet lines 350 (350A and 350Bindicated in the embodiment of FIG. 2) into the coiled tubes 410.

The amount of superheated steam in the reformer feed mixture is afunction of the nature of the carbonaceous material (i.e. the feedstock)used. In addition to steam necessary for carbonaceous feed transport,steam provides the additional hydrogen necessary to produce, from thefeedstock, suitable synthesis gas for subsequent production of liquidhydrocarbons, alcohols and/or other oxidized compounds, or othersynthesis gas conversion products therefrom. In terms of thestoichiometric ratio of carbon to hydrogen in lower alcohols such asmethanol and ethanol and C⁵⁺ hydrocarbons, the dry feedstock may have astoichiometric excess of carbon relative to hydrogen. Thus water, eithertrapped in the feedstock or in the form of superheated steam, or both,can serve to provide additional hydrogen to maximize subsequentproduction of synthesis gas conversion products. In embodiments, priorto mixing, the feedstock is relatively dry, and sufficient water isprovided by combining superheated steam with the dried feedstockmaterial in mixing apparatus 300, as discussed hereinbelow.

In embodiments, from about 0.14 kilograms (0.3 pounds) to about 0.32kilograms (0.7 pounds), from about 0.14 kg (0.3 pounds) to about 0.23 kg(0.5 pounds) or from about 0.14 kg (0.3 pounds) to about 0.18 kg (0.4pounds) of steam is added per pound of ‘dry’ feedstock comprising fromabout 4% to about 20% moisture by weight, from about 9% to about 18%moisture or from about 10% to about 20% moisture, to provide thereformer feed mixture that is introduced into the coiled tubes of thereformer. The reformer feed mixture can have a total water to feedstockratio in the range of from about 0.1 to 0.5, from about 0.2 to about0.45 or from about 0.4 to about 0.5.

Feedstock reformation carried out in the feedstock reformer isendothermic. Thus, reforming apparatus 400 comprises one or more burners404 operable to provide the necessary heat of the pyrolysis,gasification and/or reforming reaction(s) occurring within the coiledtubes 410 by combusting fuel in the presence of oxygen.

Burners 404 are desirably positioned at or near the bottom of thereformer. Burners 404 may be positioned internal or external to firebox407. In embodiments, burner(s) 404 are internal to firebox 407. Theburner(s) 404 may be distributed substantially uniformly along thediameter of vessel 407. In embodiments, the reformer has from about 1 toabout 10 burners, from about 1 to about 5 burners, or from about 1 toabout 2 burners. Oxidant utilized by the burner(s) may be provided asair, enriched air, or substantially pure oxygen. For example, in theembodiment of FIG. 2, each of the burners 404 is provided with air viaone or more air inlet lines 405 and fuel provided via one or more fuelinlet lines 406. The oxidant and fuel may be fed separately to eachburner 404 or combined prior to entry thereto. The system can furthercomprise a forced draft (FD) fan 409 configured to provide air to an airpreheater 413 configured to raise the temperature of the inlet air froma first temperature (e.g. ambient temperature) to a temperature in therange of from about −18° C. (0° F.) to about 399° C. (750° F.), fromabout 38° C. (100° F.) to about 399° C. (750° F.) or from about 316° C.(600° F.) to about 399° C. (750° F.). In embodiments, flue gas exitingsteam generation apparatus 500A (discussed further hereinbelow) isutilized to heat the air upstream of burner(s) 404. The air may bepreheated by heat transfer with a flue gas stream in steam generatorflue gas outlet line(s) 570 exiting steam generator 501A. This flue gasmay have a temperature in the range of from about 649° C. (1200° F.) toabout 1260° C. (2300° F.), from about 760° C. (1400° F.) to about 1204°C. (2200° F.) or from about 871° C. (1600° F.) to about 1149° C. (2100°F.).

Fuel is provided to the one or more burners 404 via fuel inlet line(s)406. Any fuel known in the art can be utilized. In embodiments, the fuelprovided to the reformer is selected from the group consisting ofmethane (e.g. natural gas), synthesis gas (e.g. excess synthesis gas),tailgas (e.g. Fischer-Tropsch tailgas) and combinations thereof. Asdiscussed in detail hereinbelow, in embodiments comprising tailgasrecycle line(s) 770, one or more of the burners 404 may be speciallydesigned for burning tailgas or a mixture of tailgas with at least oneother gas such as methane or synthesis gas. The amount of air combinedwith the fuel will be adjusted as known in the art based upon the fuelutilized and the desired temperature within the reformer. Inembodiments, the reformer temperature is maintained at a temperature inthe range of at least 926° C. (1700° F.), 982° C. (1800° F.), 1038° C.(1900° F.), 1093° C. (2000° F.), 1149° C. (2100° F.).

For greater energy independence of the overall system, excess synthesisgas can be made and used to run a turbine and generate electricity topower the compressors and other electrically driven devices.

The reformer comprises one or more reformer flue gas outlet lines 470for flue gas exiting the reformer. Desirably, reformer flue gas outletline(s) 470 is positioned at or near the top of the reformer. In theembodiment of FIG. 2, reformer flue gas outlet lines 470 are provided amanifold 408 fluidly connecting reformer 400A with steam generationapparatus 500A. The flue gas exiting reformer 400A can have atemperature in the range of at least 926° C. (1700° F.), 982° C. (1800°F.), 1038° C. (1900° F.), 1093° C. (2000° F.), 1149° C. (2100° F.). Thepressure of the flue gas can be in the range of from about −20 inchesH₂O to 0 inch H₂O; from about −16 inches H₂O to −2 inches H₂O; or fromabout −15 inches H₂O to −5 inches H₂O. In embodiments, the reformer isconfigured for operation at a pressure of greater than or equal to 5psig (34.5 kPa), 30 psig (206.8 kPa), 40 psig (275.8 kPa), 45 psig(310.3 kPa) or 50 psig (344.7 kPa). Operation of the reformer at higherpressures may allow a reduction in the number of compression stagesrequired upstream of the synthesis gas conversion apparatus 700 and/or areduction in required compression horsepower.

Superheated steam from line(s) 550 carries the feedstock to thereformer. In the process of heating up the feedstock upon mixingtherewith, the steam may cool to a temperature in the range of fromabout 150° F. (66° C.) to about 1000° F. (538° C.), from about 200° F.(93° C.) to about 750° F. (399° C.), or from about 300° F. (149° C.) toabout 400° F. (204° C.). In the process of heating up the feedstock uponmixing therewith, the steam may cool to a temperature of approximately204° C. (400° F.) as the reformer feed mixture approaches the reformer.In embodiments, the inlet temperature of the reformer feed mixtureentering the reformer is at a temperature of about 204° C. (400° F.).The exit temperature of the synthesis gas leaving the reformer can be inthe range of from about 870° C. (1600° F.) to about 1205° C. (2200° F.)or from about 895° C. (1650° F.) to about 930° C. (1700° F.). Inembodiments, the reformer is operated at a pressure in the range of fromabout 135 34.5 kPa (5 psig) to about 275.8 kPa (40 psig).

Within the coiled tubes of the reformer, the carbonaceous materials inthe reformer feed are anaerobically reformed with superheated steam toproduce a product process gas comprising synthesis gas (hydrogen andcarbon monoxide). The process gas can further comprise other components,for example, methane, carbon dioxide, and etc. Minor amounts of otheringredients may be formed. The reformer can comprise an external (see414A in FIG. 2) or internal (see 414B in FIG. 3) manifold configured tocombine the process gas from each of the coiled tubes 410 into one ormore reformer process gas outlet lines 480. As indicated in theembodiment of FIG. 2, outlet lines 402 associated with each of thecoiled tubes can be combined via manifold 414A to provide process gas toreformer process gas outlet line 480. In embodiments, the reformer isconfigured to provide temperature, pressure and residence timeconditions suitable to provide a process gas comprising synthesis gashaving a desired molar ratio of H₂ to CO. In embodiments, the reformeris configured to provide a synthesis gas having a H₂:CO molar ratio inthe range of from about 0.7:1 to about 2:1, from about 0.7:1 to about1.5:1 or about 1:1. In embodiments, the reformer is configured toprovide a residence time within the reformer in the range of from about0.3 s to about 3 s, from about 0.3 s to about 2 s, from about 0.3 s toabout 1 s, or from about 0.4 s to about 0.6 s.

For any given feedstock, a desired composition of the resulting processgas (i.e. the proportions of hydrogen, carbon dioxide, carbon monoxideand methane) can be provided by adjusting the contact time in thereformer, the temperature at the reformer outlet, the amount of steamintroduced with the feed, and to a lesser extent, the reformer pressure.In embodiments, the synthesis gas is to be utilized downstream for theproduction of liquid hydrocarbons via Fischer-Tropsch conversion. Inembodiments, the synthesis gas is to be utilized downstream for theproduction of liquid hydrocarbons via Fischer-Tropsch conversion with aniron-based catalyst. In such embodiments, the system may be operatedwith a reformer exit temperature in the range of from about 898° C.(1650° F.) to about 926° C. (1700° F.) and a residence or contact timethat is in the range of from about 0.3 seconds to about 2.0 seconds inthe reformer. The contact or residence time can be calculated bydividing the internal volume of the reformer by the flow rate of theprocess gas exiting the reformer.

Mixing Apparatus 300.

As indicated in FIG. 1, the biorefinery of this disclosure furthercomprises mixing apparatus 300 upstream of reformer 400. Mixingapparatus 300 is configured to combine feedstock introduced thereto viafeedstock inlet line 250 with superheated steam introduced thereto viasuperheated steam line 550. As discussed further hereinbelow, thefeedstock can be provided via feedstock handling and/or drying apparatus200 positioned upstream of mixing apparatus 300. As discussed furtherhereinbelow, superheated steam can be provided via steam generationapparatus 500 configured to utilize the heat from the reformer flue gasand/or the reformer product gas to produce superheated steam from boilerfeed water (BFW).

As depicted in the embodiment of FIG. 2, mixing apparatus 300A cancomprise one or more mixers 310 (two mixers, 310A and 310B, indicated inFIG. 2) configured to combine superheated steam with feedstock material.Feedstock can be introduced into the mixing apparatus via one or morefeedstock inlet lines 250. The feedstock comprises at least onecarbonaceous material. In embodiments, the feedstock comprises biomass.The feedstock can comprise, by way of non-limiting examples, lignite,coal, red cedar, southern pine, hardwoods such as oak, cedar, maple andash, bagasse, rice hulls, rice straw, weeds such as kennaf, sewersludge, motor oil, oil shale, creosote, pyrolysis oil such as from tirepyrolysis plants, used railroad ties, dried distiller grains, cornstalks and cobs, animal excrement, straw, or some combination thereof.The hydrogen and oxygen content for the various materials differ and,accordingly, operation of the system (e.g. amount of superheated steamcombined with the feedstock in the mixing apparatus, the reformertemperature and pressure, the reformer residence time) can be adjustedas known in the art to provide a process gas comprising synthesis gashaving a suitable molar ratio of H₂:CO for a desired subsequentsynthesis conversion application. The feedstock introduced into themixing apparatus can have an average particle size in the range of fromabout 3.9E−5 inch (0.0001 cm) to about 1 inch (2.54 cm), from about 0.01inch (0.0254 cm) to about 0.5 inch (1.27 cm) or from about 0.09 inch(0.24 cm) to about 0.2 inch (0.508 cm). In embodiments, the feedstockintroduced into the mixing apparatus has an average particle size ofless than about 1 inch (2.54 cm), less than about 0.5 inch (1.27 cm) orless than about 3/16 inch (0.48 cm). The feedstock introduced into themixing apparatus can have a moisture content in the range of from about4 weight percent to about 20 weight percent, from about 5 weight percentto about 20 weight percent, from about 10 weight percent to about 20weight percent or from about 5 weight percent to about 18 weightpercent. As discussed further hereinbelow and mentioned hereinabove, asystem of this disclosure can further comprise, upstream of the mixingapparatus and connected therewith via one or more lines 250, feedstockhandling and/or drying apparatus 200.

Within the mixing apparatus 300, feedstock is combined with superheatedsteam to provide a reformer feed mixture. In the embodiment of FIG. 2,feedstock in line 250 is divided via lines 250A and 250B and introducedinto mixers 310A and 310B respectively. Superheated steam, which may beproduced via steam generation apparatus 500 as further describedhereinbelow, is introduced via superheated steam lines 550, 550A and550B to mixing apparatus 300A. In embodiments, one or more spentcatalyst recycle lines 755 is configured to directly or indirectlyrecycle at least a portion of a catalyst/conversion product (e.g.catalyst/wax or catalyst/alcohol) stream separated from the conversionproduct within synthesis gas conversion apparatus 700 to the reformer,as discussed further hereinbelow. In embodiments, the mixing apparatusis configured to combine the feedstock in feedstock line 250 withsuperheated steam having a temperature in the range of from about 400°F. (204.4° C.) to about 1000° F. (537.8° C.), from about 600° F. (315.6°C.) to about 950° F. (510° C.) or from about 400° F. (204.4° C.) toabout 900° F. (482.2° C.) and/or a pressure in the range of from about150 psig (1034.2 kPa) to about 400 psig (2757.9 kPa), from about 200psig (1378.9 kPa) to about 375 psig (2585.5 kPa) or from about 250 psig(1723.7 kPa) to about 350 psig (2413.2 kPa). In embodiments, a system ofthis disclosure further comprises steam generation apparatus 500configured to provide superheated steam for introduction into mixingapparatus 300 as further described hereinbelow.

In the embodiment of FIG. 2, superheated steam is introduced into eachof the mixers 310A and 310B, respectively, via superheated steam lines550A and 550B. The reformer feed mixture comprising feedstock and steamis introduced into the reformer via one or more reformer inlet lines350. The feedstock/steam mixture from each mixer 310 may be introducedinto a coiled tube 410. For example, in the embodiment of FIG. 2,feedstock/steam exiting mixers 310A and 310B via lines 350A and 350B,respectively, are introduced into coiled tubes 410A and 410B,respectively. In the embodiment of FIG. 3, the feedstock/steam mixtureexiting mixing vessel 310C of system 100B is introduced via line 350,reformer feed distributor 412 and feed inlet lines 350A and 350B intocoiled tubes 410A and 410B, respectively. Other combinations of numberof mixers, manifolding of the outlets thereof, and distributors areenvisioned and not beyond the scope of this disclosure.

As indicated in FIG. 3, the mixing vessel 310C can be a cylindricalvessel having a conical bottom 320. In embodiments, superheated steam isintroduced at or near the bottom or into a conical section 320 at ornear the bottom of the mixer. Feedstock may be introduced, inembodiments, at or near the top of the mixer. In embodiments, themixture exits out the bottom of the mixing vessel.

In embodiments, the mixing vessel(s) (310A/310B/310C) are pressurevessels configured for operation at a pressure in the range of fromabout 5 psig (34.5 kPa) to about 50 psig (344.7 kPa), from about 30 psig(206.8 kPa) to about 50 psig (344.7 kPa), from about 45 psig (310.3 kPa)to about 50 psig (344.7 kPa), or configured for operation at or greaterthan about 30 psig (206.8 kPa), 45 psig (310.3 kPa) or 50 psig (344.7kPa). In embodiments, the mixing vessels are configured for operation ata temperature in the range of from about 150° F. (66° C.) to about 1000°F. (538° C.), from about 200° F. (93° C.) to about 750° F. (399° C.), orfrom about 300° F. (149° C.) to about 400° F. (204° C.).

The mixing apparatus may be configured to provide a reformer feedmixture by combining from about 0.3 pound of steam per pound offeedstock to about 0.4, 0.5, 0.6, 0.7, 0.8, 0.9 or 1 pound ofsuperheated steam per pound of feedstock. In embodiments, the mixingapparatus is configured to provide a reformer feed mixture by combiningless than or equal to about 1 pound of superheated steam per pound offeedstock.

As indicated in FIG. 3 and discussed further hereinbelow, a portion ofthe saturated steam exiting the steam generator via one or more steamgenerator steam outlet line(s) 560 can be sent via one or more line(s)560A and 560C to an excess steam condenser 516. Condensate from excesssteam condenser 516 can be combined with condensate from elsewhere inthe system (for example, with condensate in condensate outlet line 282from a dryer air preheater of feed handling and/or drying apparatus 200,as discussed further hereinbelow). Condensate can be collected fordisposal and/or recycle and reuse via line 283.

Steam Generation Apparatus 500.

The biorefinery disclosed herein further comprises steam generationapparatus 500 configured to provide superheated steam for reformingfeedstock within reformer 400/400A/400B. As depicted in the embodimentof FIG. 1, water (e.g. boiler feed water or BFW) is introduced intosteam generation apparatus 500 via one or more BFW inlet lines 580,‘hot’ reformer flue gas is introduced into steam generation apparatus500 via one or more reformer flue gas outlet lines 470, ‘hot’ productprocess gas is introduced into steam generation apparatus 500 via one ormore reformer process gas outlet lines 480, superheated steam exitssteam generation apparatus 500 via one or more superheated steam outletlines 550, saturated steam exits steam generation apparatus 500 via oneor more steam generator steam outlet lines 560, ‘cool’ flue gas exitssteam generation apparatus 500 via one or more steam generator flue gasoutlet lines 570 and ‘cool’ process gas exits steam generation apparatus500 via one or more steam generator process gas outlet lines 450.

Description of a suitable steam generation apparatus will now be madewith reference to FIG. 2. In the embodiment of FIG. 2, steam generationapparatus 500A comprises reformer flue gas and reformer effluent steamgenerator 501A and steam superheater 501B. Reformer flue gas andreformer effluent steam generator 501A is configured to producesaturated steam by heat transfer from the ‘hot’ reformer effluentprocess gas and the ‘warm’ reformer flue gas exiting steam superheater501B. Reformer effluent process gas is introduced into reformer flue gasand reformer effluent steam generator 501A via reformer process gasoutlet line(s) 480. The ‘hot’ process gas introduced into reformer fluegas and reformer effluent steam generator 501A via reformer process gasoutlet line(s) 480 may have a temperature in the range of from about870° C. (1600° F.) to about 1205° C. (2200° F.) or from about 895° C.(1650° F.) to about 930° C. (1700° F.). In embodiments, the ‘hot’process gas has a pressure in the range of from about 34.5 kPa (5 psig)to about 275 KPa (40 psig). Within reformer flue gas and reformereffluent steam generator 501A, steam is commonly generated from the fluegas and the process gas, although the two gases are not mixed. ‘Cool’reformer process gas leaves reformer flue gas and reformer effluentsteam generator 501A via steam generator process gas outlet line(s) 450.The ‘cool’ process gas exiting reformer flue gas and reformer effluentsteam generator 501A via steam generator process gas line(s) 450 mayhave a temperature in the range of from about 400° C. (752° F.) to about800° C. (1472° F.), from about 400° C. (752° F.) to about 600° C. (1112°F.) or about 400° C. (752° F.) and/or a pressure in the range of fromabout 5 psig (34.5 kPa) to about 50 psig (344.7 kPa), from about 10 psig(68.9 kPa) to about 40 psig (275.8 kPa) or from about 20 psig (137.9kPa) to about 30 psig (206.8 kPa).

Reformer flue gas is introduced into reformer flue gas and reformereffluent steam generator 501A via reformer flue gas outlet line(s) 470.The ‘hot’ flue gas introduced into reformer flue gas and reformereffluent steam generator 501A via reformer flue gas outlet line(s) 470may have a temperature in the range of from about 530° F. (276.7° C.) toabout 1500° F. (815.6° C.), from about 530° F. (276.7° C.) to about1200° F. (648.9° C.) or about 530° F. (276.7° C.) and/or a pressure inthe range of from about −20 inches H₂O to 0 inches H₂O; from about −15inches H₂O to about −5 inches H₂O; or from about −10 inches H₂O to about−5 inches H₂O. As depicted in FIG. 2, in embodiments the reformer fluegas passes through steam superheater 501B, as discussed furtherhereinbelow, prior to introduction into reformer flue gas and reformereffluent steam generator 501A. In such instances, the ‘warm’ flue gasintroduced into the reformer flue gas and reformer effluent steamgenerator 501A may have a temperature in the range of from about 1350°F. (732.2° C.) to about 2050° F. (1121.1° C.), from about 1450° F.(787.8° C.) to about 1950° F. (1065.6° C.) or from about 1350° F.(732.2° C.) to about 1850° F. (1010° C.) and/or a pressure in the rangeof from about −20 inches H₂O to 0 inch H₂O; −16 inches H₂O to −5 inchesH₂O; −15 inches H₂O to 5 inches H₂O. In embodiments, the temperature ofthe ‘warm’ flue gas is about 150 degrees less than that of the ‘hot’flue gas, i.e. the flue gas temperature drop across 501B is in the rangeof from about 130-170 degrees, from about 140-160 degrees, or about 150degrees.

‘Cool’ reformer flue gas leaves reformer flue gas and reformer effluentsteam generator 501A via steam generator flue gas outlet line(s) 570.The ‘cool’ flue gas exiting reformer flue gas and reformer effluentsteam generator 501A via steam generator flue gas outlet line(s) 570 mayhave a temperature in the range of from about 50° F. (10° C.) to about400° F. (204.4° C.), from about 200° F. (93.3° C.) to about 400° F.(204.4° C.) or about 400° F. (204.4° C.) and/or a pressure in the rangeof from about −20 inches H₂O to about 20 inches H₂O; from about −16inches to about 20 inches H₂O; or from about −15 inches H₂O to about −10inches H₂O. Induced draft (ID) fan 573 can serve to draw ‘cool’ reformerflue gas exiting reformer flue gas and reformer effluent steam generator501A via steam generator flue gas outlet line(s) 570 through airpreheater 413, discussed hereinabove. Heat transfer to the air withinair preheater 413 may provide a ‘cold’ flue gas for use elsewhere in thesystem, for example in a dryer air heater of a feed handling and/ordrying apparatus 200, as further discussed hereinbelow. The ‘cold’ fluegas passing out of air preheater 413 in line(s) 570 may have atemperature in the range of from about −18° C. (0° F.) to about 399° C.(750° F.), from about 38° C. (100° F.) to about 399° C. (750° F.) orfrom about 316° C. (600° F.) to about 399° C. (750° F.) and/or apressure in the range of from about −20 inches H₂O to about 20 inchesH₂O; from about −16 inches to about 20 inches H₂O; or from about −15inches H₂O to about −10 inches H₂O.

One or more steam generator steam outlet lines 560 carries steam (e.g.saturated steam) from reformer flue gas and reformer effluent steamgenerator 501A. A portion of the saturated steam may be directed via oneor more steam export lines 560A for export to another apparatus or useelsewhere in the system. As indicated in the embodiment of FIG. 2, allor a portion of the saturated steam produced in reformer flue gas andreformer effluent steam generator 501A can be directed to steamsuperheater 501B configured to produce superheated steam. Steamsuperheater 501B is configured to provide superheated steam at atemperature in the range of from about 400° F. (204.4° C.) to about1000° F. (537.8° C.), from about 600° F. (315.6° C.) to about 950° F.(510° C.) or from about 400° F. (204.4° C.) to about 900° F. (482.2° C.)and/or a pressure in the range of from about 150 psig (1034.2 kPa) toabout 400 psig (2757.9 kPa), from about 200 psig (1379 kPa) to about 375psig (2585.5 kPa) or from about 250 psig (1723.7 kPa) to about 350 psig(2413.2 kPa). In embodiments, steam superheater 501B operates via heattransfer from the ‘hot’ reformer flue gas in reformer flue gas outletline(s) 470. Steam superheater 501B may be configured on a manifold orheader 408 comprising reformer flue gas outlet(s) 470. As mentionedhereinabove, the ‘warm’ flue gas exiting the steam superheater may havea temperature in the range of from about 1500° F. (815.6° C.) to about2200° F. (1204.4° C.), from about 1600° F. (871.1° C.) to about 2150° F.(1176.7° C.) or from about 1600° F. (871.1° C.) to about 2100° F.(1148.9° C.) and/or a pressure in the range of from about −20 inches H₂Oto 0 inches H₂O; −16 inches H₂O to −5 inches H₂O; −15 inches H₂O to 5inches H₂O. As discussed hereinabove, superheated steam exiting steamsuperheater 501B can be introduced into the mixing apparatus 300 via oneor more superheated steam lines 550.

Reformer flue gas and reformer effluent steam generator 501A may, asknown in the art, be associated with one or more blowdown drums 515configured to purge water off and control the solids level withinreformer flue gas and reformer effluent steam generator 501A.

Description of a suitable steam generation apparatus according toanother embodiment of this disclosure will now be made with reference toFIG. 3. In the embodiment of FIG. 3, the steam generation apparatus 500Bcomprises flue gas steam generator 501A″ and reformer effluent steamgenerator 501A′. In the embodiment of FIG. 3, ‘hot’ reformer effluentprocess gas exiting reformer 400B via reformer process gas outlet lines480 passes through reformer effluent steam generator 501A′, configuredfor transfer of heat from the ‘hot’ reformer process gas to BFWintroduced thereto via BFW inlet line 580. ‘Cool’ process gas exitingreformer effluent steam generator 501A′ via steam generator process gasoutlet line 450 may have a temperature in the range of from about 752°F. (400° C.) to about 1472° F. (800° C.), from about 752° F. (400° C.)to about 1112° F. (600° C.) or about 752° F. (400° C.) and/or a pressurein the range of from about 5 psig (34.5 kPa) to about 50 psig (344.7kPa), from about 10 psig (68.9 kPa) to about 40 psig (275.8 kPa) or fromabout 20 psig (137.9 kPa) to about 30 psig (206.8 kPa).

Reformer flue gas outlet line(s) 470 may fluidly connect reformer 400Bwith steam superheater 501B′. As discussed in regard to FIG. 2, steamsuperheater 501B′ is configured to produce superheated steam having atemperature in the range of from about 400° F. (204.4° C.) to about1000° F. (537.8° C.), from about 600° F. (315.6° C.) to about 950° F.(510° C.) or from about 900° F. (482.2° C.) and/or a pressure in therange of from about 150 psig (1034.2 kPa) to about 400 psig (2757.9kPa), from about 200 psig (1379 kPa) to about 375 psig (2585.5 kPa) orfrom about 250 psig (1723.7 kPa) to about 350 psig (2413.2 kPa). One ormore superheated steam lines 550 are configured to carry the superheatedsteam from steam superheater 501B′ to mixing vessel(s) 310C. The ‘warm’flue gas exiting steam superheater 501B′ has a temperature in the rangeof from about 1350° F. (732.2° C.) to about 2050° F. (1121.1° C.), fromabout 1450° F. (787.8° C.) to about 1950° F. (1065.6° C.) or about 1850°F. (1010° C.) and/or a pressure in the range of from about −20 inchesH₂O to 0 inch H₂O; −16 inches H₂O to −5 inches H₂O; −15 inches H₂O to 5inches H₂O and passes through flue gas steam generator 501A″, configuredfor transferring heat from the ‘warm’ reformer flue gas to the steam inline 580A. One or more lines 560 are configured to carry saturated steamexiting flue gas steam generator 501A″.

One or more steam generator flue gas outlet lines 570 are configured tocarry ‘cool’ flue gas from flue gas steam generator 501A″. As mentionedhereinabove, the ‘cool’ flue gas exiting flue gas steam generator 501A″can have a temperature in the range of from about 50° F. (10° C.) toabout 400° F. (204.4° C.), from about 200° F. (93.3° C.) to about 400°F. (204.4° C.) or about 400° F. (204.4° C.) and/or a pressure in therange of from about −20 inches H₂O to about 20 inches H₂O; from about−16 inches to about 20 inches H₂O; or from about −15 inches H₂O to about−10 inches H₂O. As discussed with regard to FIG. 2, the ‘cool’ flue gasin steam generator flue gas outlet line 570 may be used to heatcombustion air in combustion air preheater 413. Combustion air preheater413 may be configured to heat air introduced thereto via FD fan 406 andone or more air inlet lines 405 from a first lower temperature (e.g.ambient temperature) to a second higher temperature in the range of fromabout from about 38° C. (100° F.) to about 399° C. (750° F.), from about316° C. (600° F.) to about 399° C. (750° F.) or about 399° C. (750° F.)for introduction into the reformer burner(s). ‘Cold’ flue gas exitingair preheater 413 may have a temperature in the range of from about −18°C. (0° F.) to about 399° C. (750° F.), from about 38° C. (100° F.) toabout 399° C. (750° F.) or from about 316° C. (600° F.) to about 399° C.(750° F.) and/or a pressure in the range of from about −20 inches H₂O toabout 20 inches H₂O; from about −16 inches to about 20 inches H₂O; orfrom about −15 inches H₂O to about −10 inches H₂O. The ‘cold’ flue gasmay be utilized elsewhere in the refinery, for example, in a dryer airheater of a feed handling and/or drying apparatus, as further discussedhereinbelow.

It will be apparent to those of skill in the art that flue gas steamgenerator 501A″ and reformer effluent steam generator 501A′ of theembodiment of FIG. 3 may be combined within a single vessel as indicatedin the embodiment of FIG. 2.

Synthesis Gas Clean-Up and Conditioning Apparatus 600.

The biorefinery disclosed herein can further comprise synthesis gascleanup and/or conditioning apparatus 600 configured to prepare thesynthesis gas for introduction into synthesis gas conversion apparatus700. As indicated in FIG. 1, synthesis gas cleanup and/or conditioningapparatus 600 is located downstream of reforming apparatus 400 and steamgeneration apparatus 500. Thus, the biorefinery is configured such thatsynthesis gas produced in the reforming apparatus 400 is, after passingthrough steam generation apparatus 500, introduced into synthesis gascleanup and/or conditioning apparatus 600.

Syngas cleanup and conditioning is a key technical barrier to thecommercialization of biomass gasification technologies and typically hasthe greatest impact on the cost of clean synthesis gas. Generally, tarreforming catalysts have not demonstrated that they can clean andcondition raw synthesis gas to meet the strict quality standardsmandated for economically feasible downstream production of productssuch as mixed alcohols and liquid hydrocarbons therefrom. The synthesisgas cleanup and conditioning apparatus disclosed herein can be utilizedto overcome some of these deficiencies.

Synthesis gas cleanup and/or conditioning apparatus 600 is configured toremove undesirables, indicated to be removed via lines 660A and 660B inFIG. 1, from the synthesis gas produced in the reformer (i.e. to‘cleanup’ the synthesis gas) and provide a synthesis gas having adesired composition for a particular downstream application (i.e. to‘condition’ the synthesis gas). For example, synthesis gas cleanupand/or conditioning apparatus 600 can be configured to remove one ormore undesirable components including, but not limited to, ash, carbondioxide, tar, methane, sulfur compounds, (excess) hydrogen and aromaticsfrom the reformer product, providing a cleaned up and conditionedsynthesis gas having a desired molar ratio of hydrogen to carbonmonoxide and acceptable levels of other components, including but notlimited to, carbon dioxide, ash, aromatics, methane, tars, and etc.

In embodiments, synthesis gas cleanup and/or conditioning apparatus 600comprises any combination of units known in the art to be suitable forcleaning and conditioning synthesis gas for downstream Fischer-Tropschproduction of liquid fuels. In embodiments, synthesis gas cleanup and/orconditioning apparatus 600 comprises one or more units selected from ashremoval apparatus, tar removal apparatus, aromatics removal units,hydrogen adjustment units, carbon dioxide removal units, andcombinations thereof. In embodiments, synthesis gas cleanup and/orconditioning apparatus 600 comprises a nickel dual fluid bed apparatusas described in U.S. patent application Ser. No. 12/691,297, which ishereby incorporated herein in its entirety for all purposes not contraryto this disclosure.

FIG. 4 is a schematic of a synthesis gas cleanup and/or conditioningapparatus 600A suitable for use in a biorefinery according toembodiments of this disclosure. As indicated in the embodiment of FIG.4, synthesis gas cleanup and/or conditioning apparatus can comprise oneor more of ash removal apparatus 605, tar removal apparatus 610,aromatics removal apparatus 620, carbon dioxide removal apparatus 630and hydrogen adjustment apparatus 640. It is to be understood that asingle apparatus or type of apparatus may be configured to remove morethan one undesirable compound. For example, an aromatics removal unitmay also serve as a tar removal unit (e.g. a TEG unit) and/or a tarremoval unit may also serve as an ash removal unit (e.g. a venturiscrubber). It is also noted that the order of the apparatus depicted inFIG. 4 may be rearranged as known in the art depending on the specificunits incorporated into the system.

In embodiments, synthesis gas cleanup and/or conditioning apparatus 600comprises ash removal apparatus 605. The ash removal apparatus 605 isconfigured to remove ash from the synthesis gas produced in thereformer. As some of the carbonaceous material used as feedstock for theproduction of synthesis gas is not carbonaceous, the ash removalapparatus may serve to remove such non-carbonaceous materials, such asphosphates and minerals therefrom. Desirably, ash removal apparatus 605reduces the level of ash in the synthesis gas to less than 12 weightpercent ash, less than 6 weight percent ash or less than 2.3 weightpercent ash. In embodiments, ash removal apparatus 605 comprises one ormore units selected from cyclones, baghouses, and scrubbers (e.g.venturi scrubbers or quench units). In embodiments, ash removalapparatus 605 comprises a first cyclone configured to separate particleshaving an average particle size of more than 1 μm, 100 μm, or 10000 μmfrom the synthesis gas. In embodiments, ash removal apparatus 605comprises a second cyclone configured to remove particles having anaverage particle size of larger 1 μm, 100 μm, or 10000 μm from thesynthesis gas exiting a first cyclone.

As depicted in FIG. 4, the synthesis gas cleanup and/or conditioningapparatus may comprise tar removal apparatus 610 and/or aromaticsremoval apparatus 620. Any tar removal apparatus known in the art may beutilized to reduce the level of tar in the synthesis gas. Desirably, tarremoval apparatus is configured to reduce the tar level in the synthesisgas to less than 1.0 weight percent tar, less than 0.1 weight percenttar, less than 0.01 weight percent or less than 0.002 weight percenttar. In embodiments, the tar level in the synthesis gas is reduced toless than 200 mg/L. In embodiments, the synthesis gas cleanup and/orconditioning apparatus comprises a venturi scrubber, for exampledownstream of one or more cyclone(s) or baghouses of an ash removalapparatus 605. The venturi scrubber may be configured to wash outwater-soluble hydrocarbons, tars, aromatics such as benzene and anyremaining ash (e.g. fines that may have escaped upstream ash separationvia ash removal apparatus 605) from the synthesis gas. The venturiscrubber may be configured for operation with a wash liquid. The washliquid may be water. Thus, a venturi scrubber may serve as ash removalapparatus 605, a tar removal apparatus 610, and/or an aromatics removalapparatus 620 (e.g. a benzene recovery unit) in a single unit. Inembodiments, the synthesis gas cleanup and/or conditioning apparatuscomprises one or more triethylene glycol units (TEG units) configuredfor solvent extraction of tars and aromatics from the synthesis gas. Inembodiments, one or more TEG units are positioned downstream of one ormore cyclones of ash removal apparatus 605 and downstream of a venturiscrubber ash removal unit 605/tar removal unit 610. The TEG unit mayserve as tar removal unit 610 and aromatics removal unit 620 in a singleapparatus. In embodiments, the TEG unit(s) removes remaining tars (e.g.heavy tars) from the synthesis gas, reducing the tar level in thesynthesis gas to less than 1.0 weight percent, less than 0.1 weightpercent, or less than 0.01 weight percent. In embodiments, the TEGunit(s) reduces the BTEX level in the synthesis gas to less than 0.5weight percent, less than 0.05 weight percent, or less than 0.005 weightpercent. In embodiments, the BTEX level is reduced to less than or about60, 50 or 45 mg/L.

As indicated in FIG. 4, synthesis gas cleanup and/or conditioningapparatus 600A can comprise carbon dioxide removal apparatus 630. Inembodiments, the carbon dioxide removal apparatus reduces the carbondioxide content of the synthesis gas to less than 10 weight percent,less than 1.0 weight percent or less than 0.1 weight percent. Anyapparatus known in the art for the removal of carbon dioxide from asynthesis gas stream may be implemented in the biorefinery of thisdisclosure. In embodiments, the synthesis gas cleanup and/orconditioning apparatus comprises an acid gas removal unit (AGRU)configured to remove carbon dioxide from synthesis gas introducedthereto. In embodiments, the synthesis gas cleanup and/or conditioningapparatus comprises an amine unit configured to remove hydrogen sulfideand carbon dioxide from synthesis gas introduced thereto.

The amine unit(s) removes carbon dioxide from the syngas. Inembodiments, a pressure swing adsorbent (PSA) unit, discussed below inconnection with hydrogen removal, could be used instead of an aminescrubber to remove the carbon dioxide. In an amine unit, the synthesisgas is scrubbed with an amine-based solvent in an absorption column. Thesolvent is regenerated in a second column thereby releasing a highpurity CO₂ product.

The carbon dioxide removal apparatus 630 serves as one point source ofcarbon dioxide sequestration provided by the disclosed biorefinery. Thecarbon dioxide removed from the synthesis gas may be sequestered andsold, for example, for use in enhanced oil recovery (EOR), as known inthe art. Sulfur compounds that may be removed in the carbon dioxideremoval apparatus (e.g. via one or more AGRU's) may be used to producevaluable commodities such as, but not limited to, fertilizer andsulfuric acid. Sequestration of carbon dioxide in this manner isenvironmentally-friendly, as it allows for a substantial reduction inthe amount of carbon dioxide, a ‘greenhouse’ gas, that is ultimatelydisposed via undesirable venting to the atmosphere.

As indicated in FIG. 4, the synthesis gas cleanup and/or conditioningapparatus can comprise hydrogen adjustment apparatus 640. Depending onthe ultimate application intended for the synthesis gas, adjustment ofthe hydrogen content in the synthesis gas may be desirable. For example,for use in Fischer-Tropsch production of liquid hydrocarbons viaFischer-Tropsch synthesis over an iron-based catalyst, it may bedesirable to remove hydrogen from the synthesis gas upstream of aFischer-Tropsch synthesis reactor in order to reduce the molar ratio ofhydrogen to carbon monoxide (e.g. to provide a hydrogen to carbonmonoxide molar ratio of about 1:1). Desirably, the reformer 400 isoperated with such a composition of feed (i.e. moisture and/or steamcontent) and at appropriate temperature, pressure, and residence timethat the synthesis gas produced therein has the desired molar ratio ofhydrogen to carbon monoxide. However, in embodiments, subsequenthydrogen adjustment may be necessary to provide a desired ratio forintroduction into subsequent Fischer-Tropsch processes.

In embodiments, hydrogen adjustment apparatus 640 is configured toincrease the molar ratio of hydrogen to carbon monoxide in the synthesisgas (i.e. to increase the hydrogen content). In such embodiments,hydrogen adjustment apparatus 640 may comprise a water gas shift reactor(WGSR) configured to produce additional hydrogen and carbon dioxide fromwater and some of the carbon monoxide in the synthesis gas, as known inthe art. In such embodiments, it may be desirable to position the WGSRupstream of the carbon dioxide removal apparatus 630 in order to allowsubsequent removal of the carbon dioxide produced in the WGSR. Inembodiments, hydrogen adjustment apparatus 640 is configured to decreasethe molar ratio of hydrogen to carbon monoxide in the synthesis gas(i.e. to decrease the hydrogen content). In such embodiments, thehydrogen adjustment apparatus 640 may comprise a hydrogen membrane orpressure swing absorber (PSA), as known in the art, configured to removehydrogen from the synthesis gas.

In embodiments, hydrogen adjustment apparatus comprises at least onePSA, as mentioned hereinabove with regard to carbon dioxide removal. Inembodiments incorporating a PSA, the synthesis gas can be compressed inone or more compressors, for example to a pressure of between 6895 KPa(1000 psi) and 16,547 KPa (2400 psi), and the resulting compressedsynthesis gas stream fed to the PSA unit(s) configured to remove aportion of the hydrogen from the synthesis gas.

Pressure swing adsorption (PSA) is an adiabatic process and is appliedfor partial hydrogen removal from synthesis gas by removing some of thehydrogen by adsorption through suitable adsorbents in fixed bedscontained in pressure vessels under high pressure. Regeneration ofadsorbents is accomplished by countercurrent depressurization and bypurging at low pressure with previously recovered hydrogen gas. Toobtain a continuous flow of product, a minimum of two adsorbers may beutilized, such that at least one adsorber is receiving feed syngas.Simultaneously, the subsequent steps of depressurization, purging andrepressurization back to the adsorption pressure are executed by theother adsorber(s). After such adsorbent regeneration andrepressurization the adsorber is switched onto adsorption duty,whereupon another adsorber is regenerated. For removing hydrogen, theadsorbent used is generally silica gel.

An alternative type of hydrogen separator which might be used toseparate a portion of the hydrogen from the synthesis gas in synthesisgas cleanup and/or conditioning apparatus 600 is a hydrogen specificpermeable membrane separator.

Synthesis Gas Conversion Apparatus 700.

The biorefinery of this disclosure further comprises synthesis gasconversion apparatus 700. As depicted in FIG. 1, synthesis gasconversion apparatus 700 is located downstream of synthesis gas cleanupand/or conditioning apparatus 600. In embodiments, synthesis gasconversion apparatus 700 is any suitable synthesis gas conversionapparatus known in the art for the production of valuable products (e.g.liquid hydrocarbons, ethanol, methanol, mixed alcohols) from synthesisgas. By way of nonlimiting examples, the synthesis gas conversionapparatus can comprise at least one selected from Fischer-Tropschreactors, alcohol synthesis reactors and microbial alcohol synthesisreactors. In embodiments, synthesis gas conversion apparatus 700comprises a Fischer-Tropsch reactor configured for the production ofliquid hydrocarbons from synthesis gas. In embodiments, theFischer-Tropsch reactor configured for the production of liquidhydrocarbons from synthesis gas is configured to operate with and/orcontains an iron-based FT catalyst or a cobalt-based FT catalyst. Inembodiments, the FT catalyst is an iron-based catalyst formed asdescribed in or having the composition of FT catalyst described in U.S.Pat. No. 5,504,118 and/or U.S. patent application Ser. Nos. 12/189,424;12/198,459; 12/207,859; 12/474,552; and/or 12/790,101, the disclosuresof each of which are hereby incorporated herein in their entirety forall purposes not contrary to this disclosure.

As indicated in FIG. 1, one or more conversion product outlet lines 750are configured for the removal of conversion product from synthesis gasconversion apparatus 700. In embodiments, the conversion productcomprises primarily liquid hydrocarbons. In embodiments, the conversionproduct comprises primarily alcohols. In embodiments, the conversionproduct comprises primarily Fischer-Tropsch hydrocarbons having five ormore carbon atoms (i.e. C⁵⁺ hydrocarbons). In embodiments, a spentcatalyst recycle line 755 is configured to directly or indirectlyrecycle at least a portion of a catalyst/conversion product (e.g.catalyst/wax or catalyst/alcohol) stream separated from the conversionproduct within synthesis gas conversion apparatus 700 to the reformer.For example, spent catalyst recycle line 755 may fluidly connectsynthesis gas conversion apparatus 700 with feedstock inlet line 250such that the catalyst/product may be combined with superheated steam inmixing apparatus 300 and subsequently introduced into the reformer. Atthe high temperatures of operation of the reformer, even long chainhydrocarbons (i.e. wax) in a conversion product can be easily convertedinto additional synthesis gas. In this manner, the hydrocarbons and/orother conversion products in the catalyst/product can be converted toadditional synthesis gas, thus improving the overall liquid yields fromthe system. Additionally, as spent catalyst is typically sent fordisposal, for example, in a landfill and since incorporation into thebiorefinery of this disclosure of such a spent catalyst recycle lineenables the spent catalyst to be separated in the ash (for example, viaash removal in the synthesis gas cleanup and/or conditioning apparatus),such recycle enables a reduction in the amount of waste material thatmust ultimately be disposed. In embodiments, the overall liquid yieldfrom a biorefinery of this disclosure comprising a Fischer-Tropschreactor configured for the production of liquid hydrocarbons is in therange of from about 0.5 to about 1.4 barrels, from about 0.6 to 2barrels, or from about 0.6 to about 1.5 barrels of conversion productper dry ton of feed material. In embodiments, the overall liquid yieldfrom a biorefinery of this disclosure is greater than or equal to about0.4, 0.5, 0.6, 0.7 0.8, 0.9, 1.0, 1.1, 1.2, 1.3 or 1.4 barrels (16.8,21, 25.2, 29.4, 33.6, 37.8, 42, 46.2, 50.4, 54.6, or 58.8 gallons) ofFischer-Tropsch conversion product per dry ton of feed material.

As indicated in FIG. 1, one or more tailgas outlet lines 760 areconfigured to remove tailgas from synthesis gas conversion apparatus700. The tailgas can comprise carbon monoxide, hydrogen, methane, carbondioxide and possibly other components. In the biorefinery of thisdisclosure, at least a portion of the tailgas produced in synthesis gasconversion apparatus 700 may beneficially be recycled via one or moretailgas recycle lines 770 to reformer 400 for use as fuel in theburner(s) thereof. In embodiments, the one or more tailgas recycle lines770 fluidly connect the one or more tailgas outlet lines 760 with one ormore of the one or more fuel inlet lines 406 feeding the one or moreburners 404 configured to provide heat to the reformer. In embodiments,the one or more tailgas recycle lines 770 fluidly connect the one ormore tailgas outlet lines 760 directly with one or more of the one ormore burners 404. In this manner, a ‘waste’ stream that is generallyconsidered of little value (i.e. tailgas) can be utilized to benefit inthe disclosed biorefinery. It is envisioned that, in embodiments, thebiorefinery further comprises a dedicated carbon dioxide removalapparatus such that carbon dioxide may be extracted from all or aportion of the tailgas exiting synthesis gas conversion apparatus 700via line 760 and/or all or a portion of the tailgas recycled via tailgasrecycle line(s) 770 and/or that a carbon dioxide removal apparatus 630of synthesis gas cleanup and/or conditioning apparatus 600 is utilizedto remove carbon dioxide therefrom. In this manner, the biorefinery andmethod of producing conversion product therefrom may be made even more‘green’, by enabling sequestration of additional carbon dioxide.Accordingly, in embodiments, tailgas recycle line(s) 770 is fluidlyconnected with an acid gas removal unit (AGRU) configured for theremoval of carbon dioxide therefrom prior to introduction of the recycletailgas into one or more burner(s) of the reformer. In embodiments, anAGRU is positioned downstream of synthesis gas conversion apparatus 700such that all or a portion of the tailgas exiting synthesis gasconversion apparatus 700 via tailgas outlet line(s) 760 can beintroduced thereto and carbon dioxide removed therefrom. In embodiments,an AGRU positioned downstream of synthesis gas conversion apparatus 700is configured to reduce the carbon dioxide content of the tailgas toless than 10 weight percent, less than 1.0 percent or less than 0.1percent. In embodiments, membrane technology is utilized to removecarbon dioxide from at least a portion of the tailgas exiting thesynthesis gas conversion apparatus 700 via tailgas outlet line(s) 760and/or from at least a portion of the tailgas utilized as fuel in one ormore burner(s) associated with the reformer. Such carbon dioxide removalfrom the tailgas can provide another point source for carbon dioxidesequestration within the disclosed biorefinery. In embodiments, at leasta portion of the carbon dioxide-reduced tailgas (containing combustiblematerial) is recycled as fuel to the reformer.

Although not specifically discussed herein, one of skill in the artwould understand that various other units may be utilized in thedisclosed biorefinery. For example, in embodiments, synthesis gascompression apparatus (e.g. synthesis gas booster compressor), as knownin the art, is positioned upstream of synthesis gas conversion apparatus700.

Feed Handling and/or Drying Apparatus 200.

A biorefinery of this disclosure may further comprise feed handlingand/or drying apparatus configured to provide feed material of a desiredaverage particle size, composition and/or moisture content to thedownstream mixing apparatus. In embodiments, the feed handling and/ordrying apparatus is substantially as disclosed in U.S. Pat. No.7,375,142, the disclosure of which is hereby incorporated herein in itsentirety for all purposes not contrary to this disclosure.

Suitable feed handling and/or drying apparatus can comprise an unloadingand tramp metal removal zone I, a comminuting zone II, a drying zoneIII, a reformer feed hopper zone IV, or some combination of two or morethereof. A feed handling and/or drying apparatus will now be describedwith reference to FIG. 5, which is a schematic of a feeding and dryingapparatus 200A according to an embodiment of this disclosure. Feedhandling and/or drying apparatus 200A comprises unloading and trampmetal removal zone I configured for unloading of feed material andremoval of undesirables therefrom. Unloading and tramp removal zone Ican comprise a truck unloading hopper 205 into which delivered feedmaterial is deposited. Truck unloading hopper 205 may be associated witha tramp metal detector 204 configured to determine the presence orabsence of undesirables such as metals in the feed material. Unloadingand tramp removal zone I can further comprise a conveyor 203 configuredto convey feed material onto a weigh belt feeder 206. A tramp metalseparator 207 is configured to remove tramp metal and other undesirablesfrom the feed material introduced thereto. Removed undesirables can beintroduced via line 208 into and stored in a bin 209 for disposal.

Comminuting zone II can be positioned downstream of unloading and trampremoval zone I, as indicated in FIG. 4, or can be downstream of anunloading zone (i.e. in the absence of a tramp removal zone).Comminuting zone II comprises apparatus configured to comminute the feedmaterial. In embodiments, the comminuting zone comprises at least onegrinder 210. A comminuting zone II may be used depending on theconsistency of the feedstock. In embodiments, the feedstock is primarilywood and/or other organic material. Grinder 210 may be used if thefeedstock is clumped together, in unusually large conglomerates, or ifthe feedstock needs to be further ground before being dried. After thefeedstock is optionally subjected to grinding, the ground material maybe passed via grinder outlet line 212 into one or more grinder dischargecyclones 220 configured to separate a larger average size fraction offeed material from a smaller sized fraction. The larger sized fractionmay be introduced via one or more grinder discharge cyclone outlet lines225 into one or more dryers 260 of dryer zone III configured to reducethe moisture content of the material fed thereto. The smaller sizedfraction from grinder discharge cyclone 220 may be passed via grinderdischarge fines outlet line 222 and grinder discharge blower 230 into adryer baghouse 240 of drying zone III, as further discussed hereinbelow.In embodiments, grinder discharge cyclone 220 is configured to providesolids having a particle size of greater than 3/16″ (0.48 cm) into dryer260 via grinder discharge cyclone outlet line 225. In embodiments,grinder discharge cyclone 220 is configured to separate solids having aparticle size of less than 3/16″ (0.48 cm) into grinder discharge finesoutlet line 222. In embodiments, grinder discharge cyclone 220 is atleast about 93, 94, 95, 96 or at least about 97 percent efficient.

Drying zone III comprises at least one dryer 260 configured to reducethe moisture content of feed material introduced therein. In theembodiment of FIG. 5, drying zone III comprises dryer 260, dryer airheater 280, dryer cyclone 265, dryer baghouse 240, accumulator 245,dryer exhaust fan 241 and dryer stack 246. Various embodiments maycomprise any combination of these components. Within drying zone III,the feedstock is dried to a moisture content in the range of from about5% to about 20%, from about 5% to about 15% or from about 9% to about15%. The flue gas and air fed into dryer 260 mixes with comminutedfeedstock to dry it, purge it and heat it for further processing.

An air supply fan 261 is configured to introduce air via line 262 andreformer flue gas (e.g. ‘cold’ reformer flue gas from air preheater 413)via line 570 into dryer air heater 280. The flue gas may be addedupstream of dryer air preheater 280 to prevent temperatures above 400°F. (204.4° C.) to the inlet of dryer 260, preventing fire therein. Asmentioned hereinabove, the ‘cold’ flue gas may have a temperature in therange of from about −18° C. (0° F.) to about 399° C. (750° F.), fromabout 38° C. (100° F.) to about 399° C. (750° F.) or from about 316° C.(600° F.) to about 399° C. (750° F.) and/or a pressure in the range offrom about −20 inches H₂O to about 20 inches H₂O; from about −16 inchesto about 20 inches H₂O; or from about −15 inches H₂O to about −10 inchesH₂O. In embodiments, the flue gas introduced via line 570 comprisesabout 80% nitrogen and 20% CO₂.

A portion of the effluent steam from reformer effluent and reformer fluegas steam generator 501A or from flue gas steam generator 501A″ can beintroduced via line 560A or 560D into dryer air preheater 280. The steamintroduced into dryer air preheater 280 may have a temperature in therange of from about 150° F. (65.6° C.) to about 500° F. (260° C.), fromabout 250° F. (121.1° C.) to about 450° F. (232.2° C.) or from about300° F. (148.9° C.) to about 400° F. (204.4° C.) and/or a pressure inthe range of from about 70 psig (482.6 kPa) to about 300 psig (2068.4kPa), from about 150 psig (1034.2 kPa) to about 300 psig (2068.4 kPa) orfrom about 250 psig (1723.7 kPa) to about 300 psig (2068.4 kPa).Condensate outlet line 282 is configured for removal of condensate fromair dryer 280. The pressure of the condensate may be reduced downstreamof the air dryer 280 and the condensate combined as indicated in FIG. 3with condensate from excess steam condenser 516. Heated air exitingdryer air heater 280 via heated air line 284 may have a temperature inthe range of from about −18° C. (0° F.) to about 204° C. (400° F.), fromabout −18° C. (0° F.) to about 149° C. (300° F.) or from about −18° C.(0° F.) to about 93.3° C. (200° F.). Desirably, the heated airtemperature does not exceed 400° F.

Heated air line 284 fluidly connects dryer air heater 280 with dryer260. Drying zone III may further comprise a heated air distributor 286configured to divide heated air line 284 into a plurality of heated airdryer inlet lines. For example, in the embodiment of FIG. 5, distributor286 divides the flow of air from heated air line 284 into three heatedair dryer inlet lines 284A-284C. Air passing through dryer 260 maycomprise entrained feed material. Accordingly, drying zone III cancomprise one or more dryer cyclones 265 configured to separate solidsfrom the air exiting dryer 260. In the embodiment of FIG. 5, air exitingdryer 260 via dryer vent lines 286A-286C is combined via air manifold287 into dryer vent line 281 which is fed into dryer cyclone 265. It isto be noted that, although three air inlet and air outlet (vent) linesare shown in the embodiment of FIG. 5, any number of air inlet lines andoutlet lines may be utilized. Additionally, the number of air inletlines to dryer 260 need not be equal to the number of air outlet or ventlines.

Dryer cyclone 265 is configured to remove solids from the vent gasexiting dryer 260. Air and any fines entrained therein exit dryercyclone 265 via dryer cyclone fines outlet line 266, while solids exitdryer cyclone 265 via dryer cyclone solids outlet line 267. Line 267 maybe fluidly connected with reformer feed hopper inlet line 276. Dryercyclone fines outlet line 266 may be configured to introduce air andentrained fines into dryer baghouse 240 along with fines introducedthereto from grinder discharge cyclone 220, grinder discharge cycloneoutlet line 222, grinder discharge blower 230 and/or grinder dischargeblower outlet line 231. In embodiments, dryer cyclone 265 is configuredto provide solids having a particle size of greater than 3/32″ (2.5 mm)or greater than 3/16″ (4.8 mm) into dryer cyclone solids outlet line267. In embodiments, dryer cyclone 265 is configured to separate solidshaving a particle size of less than 3/16″ into dryer cyclone finesoutlet line 266. In embodiments, dryer cyclone 265 has an efficiency ofat least 85, 90, 92, 95, 96, 97, or 98 percent.

One or more dryer baghouses 240 are configured to remove solids from theair introduced thereto. One or more dryer baghouse solids outlet lines243 are configured to introduce solids separated within dryer baghouse240 into reformer feed hopper cyclone inlet line 276 of reformer feedhopper zone IV, further discussed hereinbelow. In embodiments, dryerbaghouse 240 is configured to provide solids having a particle size ofgreater than 20, 15, 10 or 5 μm into dryer baghouse solids outlet line243. In embodiments, dryer baghouse 240 is configured to separate solidshaving a particle size of less than 10 μm into dryer baghouse finesoutlet line 244.

One or more dryer baghouse fines outlet lines 244 are configured tointroduce gas from dryer baghouse 240 into dryer stack 246, optionallyvia dryer exhaust fan 241 and line 247. A line 251 may introduce airinto an accumulator 245 prior to introduction into dryer baghouse(s)240.

Feed handling and/or drying apparatus 200A can further comprise areformer feed hopper zone IV. The reformer feed hopper zone IV comprisesat least one reformer feed hopper and a feeder configured for feedingfeed material into mixing apparatus 300. In the embodiment of FIG. 5,reformer feed hopper zone IV comprises reformer feed hopper 295 andmixing vessel rotary feeder 297. Reformer feed hopper zone IV canfurther comprise a surge hopper 270, a reformer feed hopper blower 275and a reformer feed hopper cyclone 290, as indicated in the embodimentof FIG. 5. One or more dried feed lines 294 are configured to introducedried feed material from one or more dryers 260 of dryer zone III intoreformer feed hopper zone IV. The feed material may be introduced into asurge hopper 270, configured for storage of surplus dried feed materialand supply therefrom to reformer feed hopper 295. A reformer feed hopperblower 275 may be incorporated into zone IV for pushing dried feedmaterial and/or separated solids introduced into reformer feed hoppercyclone inlet line 276 from dryer(s) 260 and/or surge hopper(s) 270 vialine(s) 271, from dryer cyclone(s) 265 via dryer cyclone solids outletline(s) 267, from dryer baghouse(s) 240 via dryer baghouse solids outletline(s) 243 into reformer feed hopper cyclone(s) 290. In alternativeembodiments, the material in reformer feed hopper inlet line(s) 276 isintroduced directly into reformer feed hopper 295. Reformer feed hoppercyclone 290 is configured to separate fines from material introducedtherein. In embodiments, a reformer feed hopper cyclone outlet line 292is configured to introduce fines separated within reformer feed hoppercyclone 290 into dryer baghouse 240, optionally via grinder dischargeblower outlet line 231 as indicated in the embodiment of FIG. 5. Inembodiments, reformer feed hopper cyclone 290 is configured to providesolids having an average particle size in the range of from about 3.9E−5inch (0.0001 cm) to about 1 inch (2.54 cm), from about 0.01 inch (0.0254cm) to about 0.5 inch (1.27 cm) or from about 0.09 inch (0.24 cm) toabout 0.2 inch (0.51 cm) into reformer feed hopper 295. In embodiments,the feed material in reformer feed hopper 295 is of a size allowing itto pass through a 4.8 millimeter ( 3/16 inch) screen. In embodiments,reformer feed hopper cyclone 290 is configured to separate solids havinga particle size of less than 3/16″ (0.48 cm) into reformer feed hoppercyclone fines outlet line 292. Feed material is introduced into reformerfeed hopper 295 via reformer feed hopper inlet line 276 and optionallyreformer feed hopper cyclone 290. In embodiments, reformer feed hopper295 is a cylindrical vessel having a conical bottom. In embodiments,reformer feed hopper cyclone 295 provides an efficiency of at least 80,85, 90, 92, 95, 96, 97 or 98 percent.

Mixing vessel rotary feeder 297 is configured to introduce feed materialfrom reformer feed hopper 295 into mixing apparatus 300. As needed, feedmaterial is fed from reformer feed hopper 295 and rotary feeder 297 intomixing apparatus 300. Rotary feeder 297 may be substantially asdescribed in U.S. Pat. No. 7,375,142. Feed material exits reformer feedhopper 295 via feed hopper outlet line 296, which fluidly connectsreformer feed hopper 295 with mixing vessel rotary feeder 297.

In embodiments, one or more purge lines 291 is configured to introducepurge gas (e.g. flue gas or plant air) for purge into and push feedmaterial through reformer feed hopper 295. In embodiments, the purge gasis flue gas comprising about 80% nitrogen and about 20% carbon dioxide,helping to insure that the reformation process in reformer 400 will becarried out anaerobically. Reformer feed hopper 295 may also include avent for venting flue gas. From reformer feed hopper 295, feedstocksettles into feed hopper outlet line(s) 296, which extends from thebottom of reformer feed hopper 295. The feedstock is metered by rotaryvalve 297 into feedstock inlet line 250, along which it is entrainedwith steam under pressure entering from superheated steam line 550 ofmixing apparatus 300. To keep feedstock flowing into the stream ofsteam, and in order to counter steam back pressure in line 250, a supplyof gas is moved through rotary feeder purge gas inlet line 288 via acompressor to an inlet just below valve 297. To prevent the pressure infeedstock inlet line 250 from blowing feedstock back into rotary valve297, some of the gas is also split off from rotary feeder purge gasinlet line 288 and fed to an inlet of mixing vessel rotary feeder 297.Rotary feeder 297 includes a central rotor having a plurality of vaneswhich divide the interior of valve 297 into separate compartments.Opposite the inlet on rotary valve 297, is an outlet pressure vent line289. As the rotor of valve 297 rotates, the compartment formed by thevanes at the top fill with feedstock. That filled compartment is thenrotated until it opens to the inlet, where it is pressurized withincoming gas. As the rotor rotates further, the feedstock filled andpressurized chamber opens into reformer feedstock inlet line 250. Sincethe pressure in the rotor chamber is equalized with the pressure in line250, the feedstock falls down into feedstock inlet line 250. As thevalve rotor continues on its journey, it is eventually vented throughoutlet pressure vent line 289, such that when the chamber again reachesfeed hopper outlet line 296, it is depressurized and will not vent backup into feed hopper outlet line 296. After feedstock has moved throughrotary feeder valve 297 and into feedstock line 250, it feeds by gravityinto a mixing chamber or position along mixing apparatus feedstock inletline 250 where the feedstock is mixed with superheated steam (e.g. steamhaving a temperature of about 510° C. (950° F.)) from superheated steamline 550.

II. Method of Producing Synthesis Gas Conversion Product.

Also disclosed herein is a method of producing synthesis gas conversionproduct. In embodiments, the conversion product comprises primarilyliquid Fischer-Tropsch hydrocarbons. In embodiments, the conversionproduct comprises primarily alcohols and/or other oxidized compounds.

The basic steps in the method of producing synthesis gas conversionproduct according to this disclosure are depicted in the flow diagram ofFIG. 6. As indicated in FIG. 6, a method of producing synthesis gasconversion product 801 comprises producing synthesis gas via reformingof carbonaceous material at 800, cleaning up and/or conditioning thesynthesis gas at 900, converting the synthesis gas into product at 1000and recycling 1100 at least one component from converting at 1000 forreuse in the producing of synthesis gas at 800.

Producing Synthesis Gas 800.

The method of producing synthesis gas conversion products according tothis disclosure comprises producing synthesis gas at 800. FIG. 7 is aflow diagram depicting a method 800A of producing synthesis gasaccording to an embodiment of this disclosure. Method 800A comprisespreparing carbonaceous feedstock at 810, preparing reformer feed at 820and reforming the reformer feed at 830. Preparing carbonaceous feedmaterial 810 comprises comminuting and/or drying a suitable carbonaceousfeed material. In embodiments, the source of the carbonaceous feedstockcomprises biomass. In embodiments, the carbonaceous feedstock comprisesat least one component that is or that is derived from lignite, coal,red cedar, southern pine, hardwoods such as oak, cedar, maple and ash,bagasse, rice hulls, rice straw, weeds such as kennaf, sewer sludge,motor oil, oil shale, creosote, pyrolysis oil such as from tirepyrolysis plants, used railroad ties, dried distiller grains, cornstalks and cobs, animal excrement, straw, and combinations thereof.

Preparing Carbonaceous Feedstock 810.

In embodiments, preparing the carbonaceous feedstock 810 comprisessizing (comminuting) at least one carbonaceous feedstock such that it isof a desirable size for effective reforming. In embodiments, preparingthe carbonaceous feedstock comprises reducing the average particle sizeof the feedstock to less than about ⅝^(th) inch (15.9 mm), ½ inch (12.7mm), or less than about 3/16^(th) of an inch (4.8 mm). The carbonaceousfeedstock may be sized by any methods known in the art. In embodiments,a carbonaceous material is sized by introducing it into one or moregrinders 210, as discussed above with reference to FIG. 5.

In embodiments, preparing the carbonaceous feed material comprisesdrying the carbonaceous feedstock to a moisture content in the range offrom about 4 weight percent to about 20 weight percent, from about 5weight percent to about 20 weight percent, from about 10 weight percentto about 20 weight percent or from about 5 weight percent to about 18weight percent. In embodiments, preparing the carbonaceous feedstockcomprises drying the carbonaceous feedstock to a moisture content ofless than about 25, 20, 15, 10 or 9 weight percent. The carbonaceousfeedstock may be dried by any methods known in the art. In embodiments,a carbonaceous feedstock is dried by introducing it into one or moredryers 260, as discussed above with reference to FIG. 5. In embodiments,ground carbonaceous material exiting grinder 210 is introduced into agrinder discharge cyclone 220. Within grinder discharge cyclone 220, astream of larger sized particles is separated via grinder dischargecyclone outlet line 225 from a stream of smaller sized particles ingrinder discharge fines outlet line 222. A grinder discharge blower 230may introduce the smaller particles separated in grinder dischargecyclone 220 into one or more dryer baghouse(s) 240. The larger particlesexiting grinder discharge cyclone 220 via grinder discharge cycloneoutlet line 225 are introduced into dryer 260.

In embodiments, air supplied via air supply fan 261 and line 262 iscombined with flue gas in line 570 and introduced into dryer air heater280. The flue gas utilized here may be produced during reforming of thecarbonaceous material discussed below. Heat transfer with steamintroduced into the dryer air heater via steam inlet line 560A/560Dproduces heated air in heated air line 284 and condensate in condensateoutlet line 282. As discussed hereinabove, the steam utilized in dryerair heater 280 may be produced via heat transfer with the hot reformerprocess gas effluent and/or the ‘warm’ flue gas effluent, as discussedfurther hereinbelow.

Heated air in heated air line 284 may be divided by a heated airdistributor or divider 286 into a plurality of heated air inlet lines284A-284C. Within dryer 260, the comminuted carbonaceous material isdried to a desired moisture content, as mentioned hereinabove. Dryereffluent comprising air and fines is introduced via dryer vent line 281into dryer cyclone 265. Dried carbonaceous material exits dryer 260 viaone or more dried feed lines 294 and surge hopper 270. Air from reformerfeed hopper blower 275 may push comminuted and dried feed material fromdryer 260 and surge hopper 270 along reformer feed hopper inlet line 276into reformer feed hopper cyclone 290. Solids removed from dryer cyclone265 and dryer baghouse 240 may be introduced into reformer feed hopperinlet line 276, as indicated in FIG. 5.

Gas exiting dryer cyclone 265 may be combined in grinder dischargeblower outlet line 231 via dryer cyclone fines outlet line 266 with gasexiting grinder discharge blower 220 and gas exiting reformer feedhopper cyclone 290 via line 292 and introduced into dryer baghouse 240.Gases exiting dryer baghouse via dryer baghouse fines outlet line 244may pass via dryer exhaust fan 241 and line 247 to dryer stack 246.

Dried carbonaceous materials exit reformer feed hopper cyclone 290 andenter reformer feed hopper 295. Carbonaceous material from reformer feedhopper 295 is introduced via mixing vessel rotary feeder 297 andfeedstock line 250 into one or more mixing vessels of mixing apparatus300.

Preparing Reformer Feed 820.

As discussed above, producing synthesis gas via reforming ofcarbonaceous material 800 further comprises preparing reformer feed 820.A suitable reformer feed may be formed via combination of superheatedsteam and comminuted and dried carbonaceous material via any methodsknown in the art. In embodiments, spent catalyst comprising spentcatalyst and associated synthesis gas conversion product is combinedwith the carbonaceous material prior to or along with combination withsuperheated steam. In embodiments, preparing reformer feed comprisesintroducing the comminuted and dried carbonaceous feed material andsuperheated steam into one or more mixing vessels as describedhereinabove.

With reference to FIG. 2, preparing reformer feed material can compriseintroducing comminuted and dried feed material via lines 250, 250A and250B into mixing apparatus 300A. Spent catalyst/conversion product maybe combined with the carbonaceous material via line 755. In alternativeembodiments, spent catalyst/conversion product is introduced directlyinto the mixing vessel(s). Superheated steam from steam superheater 501Bis introduced via superheated steam lines 550, 550A and 550B into mixers310A and 310B, respectively.

With reference to FIG. 3, preparing reformer feed can compriseintroducing comminuted and dried feed material via feedstock inlet line250 into mixing apparatus 300B. Superheated steam from steam superheater501W is introduced via superheated steam line 550 into mixer 310C.

As mentioned hereinabove, within the mixing apparatus, superheated steamand carbonaceous material are combined to provide a reformer feedmixture comprising from about 0.14 kilograms (0.3 pounds) to about 0.032kilograms (0.7 pounds), from about 0.14 kg (0.3 pounds) to about 0.23 kg(0.5 pounds) or from about 0.14 kg (0.3 pounds) to about 0.18 kg (0.4pounds) of steam per pound of ‘dry’ feedstock comprising from about 4%to about 20% moisture by weight, from about 9% to about 18% moisture orfrom about 10% to about 20% moisture, to provide the reformer feedmixture that is introduced into the coiled tubes of the reformer. Inembodiments, the reformer feed comprises from about 0.01 wt % to about20 wt %, from about 0.05 wt % to about 10 wt %, or from 1 wt % to about5 wt % weight percent spent catalyst/product (e.g. cat/wax). Thereformer feed may have a temperature in the range of from about 150° F.(66° C.) to about 1000° F. (538° C.), from about 200° F. (93° C.) toabout 750° F. (399° C.), or from about 300° F. (149° C.) to about 400°F. (204° C.). In embodiments, the reformer feed has a pressure of atleast or about in the range of from about 34.5 kPa (5 psig) to about 275kPa (40 psig).

The superheated steam utilized in the reformer feed mixers may beproduced by heat exchange with the reformer flue gas effluent and/or thereformer process gas effluent. With reference to FIG. 2, BFW may beintroduced via BFW inlet line(s) 580 into reformer effluent and reformerflue gas steam generator 501A. Within reformer effluent and reformerflue gas steam generator 501A, heat transfer between the hot gas (‘warm’reformer flue gas passing through steam superheater 501B and ‘hot’reformer process gas effluent) and the BFW may produce steam (in steamoutlet line 560) having a temperature in the range of from about 300° F.(148.9° C.) to about 500° F. (260° C.), from about 350° F. (176.7° C.)to about 500° F. (260° C.) or from about 350° F. (176.7° C.) to about500° F. (260° C.) and a pressure in the range of from about 200 psig(1379 kPa) to about 300 psig (2068.4 kPa), from about 250 psig (1723.7kPa) to about 300 psig (2068.4 kPa), or from about 275 psig (1896.1 kPa)to about 300 psig (2068.4 kPa). Steam exiting reformer effluent andreformer flue gas steam generator 501A via steam generator steam outletline 560 may be divided, with a portion entering steam superheater 501Bvia line 560B and another portion exported via line 560A. Within steamsuperheater 501B, heat transfer between ‘hot’ reformer flue gas andsteam produces superheated steam having a temperature in the range offrom about 400° F. (204.4° C.) to about 1000° F. (537.8° C.), from about600° F. (315.6° C.) to about 950° F. (510° C.) or from about 400° F.(204.4° C.) to about 900° F. (482.2° C.) and/or a pressure in the rangeof from about 150 psig (1034.2 kPa) to about 400 psig (2757.9 kPa), fromabout 200 psig (1379 kPa) to about 375 psig (2585.5 kPa) or from about250 psig (1723.7 kPa) to about 350 psig (2413.2 kPa). The superheatedsteam exiting steam superheater 501B is introduced into reformer feedmixing vessels 310A/310B via lines 550 and 550A/550B.

With reference to FIG. 3, BFW may be introduced via BFW inlet line 580into reformer effluent steam generator 501A′. Within reformer effluentsteam generator 501A′, heat transfer between the hot process gaseffluent and the BFW may produce steam. Steam exiting reformer effluentsteam generator 501A′ via line 580A may be introduced into flue gassteam generator 501A″. Within flue gas steam generator 501A″, heattransfer between ‘warm’ reformer flue gas and steam produces saturatedsteam (exiting via steam generator steam outlet line 560) having atemperature in the range of from about 300° F. (148.9° C.) to about 500°F. (260° C.), from about 350° F. (176.7° C.) to about 500° F. (260° C.)or from about 350° F. (176.7° C.) to about 500° F. (260° C.) and apressure in the range of from about 200 psig (1379 kPa) to about 300psig (2068.4 kPa), from about 250 psig (1723.7 kPa) to about 300 psig(2068.4 kPa), or from about 275 psig (1896.1 kPa) to about 300 psig(2068.4 kPa).

Reformer flue gas exiting the reformer via reformer flue gas outlet line470 passes through steam superheater 501B′, wherein the temperature ofthe ‘hot’ flue gas is reduced to a temperature in the range of fromabout 530° F. (276.7° C.) to about 1500° F. (815.6° C.), from about 530°F. (276.7° C.) to about 1200° F. (648.9° C.), or about 530° F. (276.7°C.), and/or a pressure in the range of from about −20 inches H₂O to 0inch H₂O; from about −15 inch H₂O to about −5 inch H₂O; or from about−10 inches H₂O to about −5 inches H₂O, and superheated steam isproduced. The superheated steam may have a temperature in the range offrom about 400° F. (204.4° C.) to about 1000° F. (537.8° C.), from about600° F. (315.6° C.) to about 950° F. (510° C.), or from about 400° F.(204.4° C.) to about 900° F. (482.2° C.), and/or a pressure in the rangeof from about 150 psig (1034.2 kPa) to about 400 psig (2757.9 kPa), fromabout 200 psig (1379 kPa) to about 375 psig (2585.5 kPa), or from about250 psig (1723.7 kPa) to about 350 psig (2413.2 kPa). The superheatedsteam exiting steam superheater 501W is introduced into reformer feedmixing vessel 310C via line 550.

Reforming Reformer Feed 830.

As discussed above, producing synthesis gas via reforming ofcarbonaceous material 800 further comprises reforming reformer feed at830. In embodiments, reforming the reformer feed 830 comprisesconverting the reformer feed into synthesis gas via introduction into areformer as described above. Reforming of the synthesis gas will now bedescribed with reference to FIGS. 2 and 3. Reformer feed is introducedinto the reformer via one or more reformer feed inlet lines 350. Inembodiments, a distributor 412 distributes the reformer feed evenlyamong a plurality of coiled tubes 410. Within the coiled tubes,reforming of the carbonaceous feedstock produces synthesis gas. Inembodiments, the temperature of the reformer (e.g. reformer effluent) ismaintained in the range of up to or about 926° C. (1700° F.), 982° C.(1800° F.), 1038° C. (1900° F.), 1093° C. (2000° F.), 1149° C. (2100°F.). In embodiments, the pressure of the reformer is maintained in therange of from about 0 psig (0 kPa) to about 100 psig (689.5 kPa), fromabout 2 psig (13.8 kPa) to about 60 psig (413.7 kPa) or from about 5psig (34.5 kPa) to about 50 psig (344.7 kPa). In embodiments, thereformer pressure is maintained at a pressure of equal to or greaterthan about 2 psig (13.8 kPa), about 5 psig (34.5 kPa), or about 50 psig(344.7 kPa).

The heat needed to maintain the desired reformer temperature is providedto the endothermic reforming process by the combustion of fuel in one ormore burners 404. Air for the combustion may be heated in air preheater413 prior to burning with the fuel in burners 404. The fuel combusted inthe burner(s) 404 may be selected from tailgas (e.g. Fischer-Tropschtailgas), synthesis gas, methane (e.g. natural gas), and combinationsthereof. Desirably, at least a portion of the fuel combusted in at leastone of the burner(s) 404 comprises tailgas recycled from the step ofconverting the synthesis gas into product 1000. At least one of theburner(s) 404 may be specially designed for the combustion of tailgas orfor the combustion of tailgas in combination with another gas, forexample in combination with a gas selected from synthesis gas andmethane (e.g. natural gas). In embodiments, recycle tailgas in line(s)770 is introduced into one or more burner(s) 404 by introduction intoone or more of the fuel lines 406 or via another fuel inlet line(s).

Cleaning Up and/or Conditioning Synthesis Gas 900.

The method of producing synthesis gas conversion products according tothis disclosure further comprises cleaning up and/or conditioning thesynthesis gas at 900. Cleaning up and/or conditioning the synthesis gascomprises removing one or more components from the synthesis gas.Cleaning up and/or conditioning the synthesis gas can comprise removingone or more components selected from the group consisting of ash, tar,aromatics, carbon dioxide, hydrogen sulfide, carbon monoxide andhydrogen from the synthesis gas. In embodiments, the synthesis gas isintroduced into one or more cyclones and/or baghouses for the removal ofash. The ash may be reduced to a level of less than 10 weight percent,less than about 1.0 weight percent or less than about 0.1 weightpercent. In embodiments, tar is removed from the synthesis gas. Inembodiments, the synthesis gas is introduced into a venturi scrubberthat serves to remove water-soluble hydrocarbons, tar, ash fines thatescaped the ash removal apparatus and/or benzene from the synthesis gas.

In embodiments, aromatics are removed from the synthesis gas. Inembodiments, ash and/or tar-reduced synthesis gas is introduced into aTEG unit whereby solvent extraction reduces the level of tars and/oraromatics therein.

In embodiments, conditioning comprises removing carbon dioxide from thesynthesis gas. In embodiments, an acid gas removal unit is utilized toremove carbon dioxide and hydrogen sulfide from the synthesis gas. TheAGRU may be downstream of a TEG unit.

In embodiments, conditioning comprises subjecting the synthesis gas towater gas shift by reaction of carbon monoxide in the synthesis gas withwater to produce carbon dioxide and additional hydrogen, thus reducingthe carbon monoxide content of the synthesis gas. In embodiments, watergas shift provides a synthesis gas having a mole ratio of hydrogen tocarbon monoxide in the range of from about 0.5:1 to about 1.1:1, fromabout 0.7:1 to about 1.1:1, or about 1:1.

In embodiments for which the synthesis gas will be used at step 1000 forthe production of liquid Fischer-Tropsch hydrocarbons, the cleaned-upand conditioned synthesis gas may have a molar ratio of hydrogen tocarbon monoxide in the range of from about 0.5:1 to about 1.1:1, fromabout 0.7:1 to about 1.1:1, or about 1:1, a tar content of less thanabout 200 ppm (e.g. ˜95% removal), a carbon dioxide content of less thanabout 10 weight % (e.g. ˜95 percent removal), a sulfur content of lessthan about 1 ppm, a BTEX content of less than about 50 ppm (e.g. ˜95%removal), an ash content of less than about 0.1 weight percent, or acombination thereof.

Converting Synthesis Gas into Product 1000.

The method of producing synthesis gas conversion product according tothis disclosure further comprises converting the synthesis gas intoproduct at 1000. Desirably, the conversion produces, in addition tovaluable primary products, a tailgas suitable for recycle to the step ofproducing synthesis gas via reforming 800, as further discussedhereinbelow. In embodiments, the tailgas comprises one or more gasselected from carbon dioxide, methane, hydrogen (e.g. unreactedhydrogen) and carbon monoxide (e.g. unreacted carbon monoxide).

The cleaned and/or conditioned synthesis gas is converted into valuableproducts at 1000. This conversion may be referred to herein asFischer-Tropsch conversion. FIG. 8 is a flow diagram of a method 1000Afor converting synthesis gas to product, according to an embodiment ofthis disclosure. Method 1000A comprises subjecting the synthesis gas toreaction conditions whereby conversion products are formed 1010 andseparating tailgas and/or spent catalyst/product from the conversionproduct(s) 1020.

Converting synthesis gas into product comprises subjecting the synthesisgas to reaction conditions whereby conversion products are formed. Inembodiments, the synthesis gas is converted into products consistingprimarily of alcohols and/or other oxidized compounds. Suitable reactionconditions including temperatures, pressures and catalysts for suchconversion are known in the art.

In embodiments, the synthesis gas is converted into products consistingprimarily of liquid hydrocarbons. In embodiments, the synthesis gas isconverted into products consisting primarily of C⁵⁺ hydrocarbons. Insuch embodiments, the synthesis gas may be compressed, as needed, andintroduced into one or more Fischer-Tropsch reactors configured for theproduction of liquid hydrocarbons. In such embodiments, subjecting thesynthesis gas to reaction conditions whereby conversion products areformed can comprise contacting the synthesis gas with an FT catalystthat promotes the FT synthesis reactions at suitable temperatures andpressures as known in the art.

In embodiments, the catalyst comprises at least one catalytically activemetal or oxide thereof. In embodiments, the catalyst further comprises acatalyst support. In embodiments, the catalyst further comprises atleast one promoter. The catalytically active metal may be selected fromthe group consisting of Co, Fe, Ni, Ru, Re, Os, and combinations of twoor more thereof. The support material may comprise alumina, zirconia,silica, aluminum fluoride, fluorided alumina, bentonite, ceria, zincoxide, silica-alumina, silicon carbide, a molecular sieve, or acombination of two or more thereof. The support material may comprise arefractory oxide. The promoter may be selected from Group IA, IIA, IIIBor IVB metals and oxides thereof, lanthanide metals and metal oxides,and actinide metals and metal oxides. In embodiments, the promoter isselected from the group consisting of Li, B, Na, K, Rb, Cs, Mg, Ca, Sr,Ba, Sc, Y, La, Ac, Ti, Zr, La, Ac, Ce and Th, oxides thereof, andmixtures of two or more thereof. Suitable catalysts may be as disclosedin U.S. Pat. Nos. 4,585,798; 5,036,032; 5,733,839; 6,075,062; 6,136,868;6,262,131; 6,353,035; 6,368,997; 6,476,085; 6,451,864; 6,490,880;6,648,662; 6,537,945; 6,558,634; and U.S. Patent App. Pub. No.2003/0105171; these patents and patent publications being incorporatedherein by reference for their disclosures of Fischer-Tropsch catalystsand methods for preparing such catalysts.

The FT catalyst can be any suitable catalyst known in the art. Inembodiments, the FT catalyst is an iron-based catalyst formed asdescribed in or having the composition of an FT catalyst described inU.S. Pat. No. 5,508,118 and/or U.S. patent application Ser. Nos.12/189,424; 12/198,459; 12/207,859; 12/474,552; and/or 12/790,101, thedisclosures of each of which are hereby incorporated herein in theirentirety for all purposes not contrary to this disclosure.

In embodiments, the FT catalyst is an iron-based catalyst comprisingiron, copper and potassium. The catalyst may have a weight ratio of100Fe:1Cu:1K (wt %:wt %:wt %), wherein the iron in the catalystcomprises a maghemite:hematite weight ratio in the range of about 1%:99%to about 70%:30%. The iron catalyst can comprise a maghemite to hematiteratio of about 10%:90%. The catalyst can have a particle sizedistribution in the range of 10 μm-100 μm. The catalyst can exhibit aBET surface area in the range of from about 45 m²/g to about 150 m²/g orfrom about 45 m²/g to about 65 m²/g. The catalyst can have a mean porediameter in the range of from about 45 Å to about 120 Å or from about 75Å to about 120 Å. The catalyst can have a mean pore volume in the rangeof from about 0.2 cc/g to about 0.6 cc/g or from about 0.20 cc/g toabout 0.24 cc/g. The catalyst can have a mean crystallite size in therange of from about 15 nm to about 40 nmor from about 25 nm to about 29nm.

Depending on the preselected alpha, i.e., the polymerization probabilitydesired, a precipitated iron catalyst may have a weight ratio ofpotassium (e.g., as carbonate) to iron in the range of from about 0.005and about 0.015, in the range of from 0.0075 to 0.0125, or about 0.010.Larger amounts of alkali metal promoter (e.g., potassium) may cause theproduct distribution to shift toward the longer-chain molecules, whilesmall amounts of alkali metal may result in a predominantly gaseoushydrocarbon product.

The weight ratio of copper to iron in the iron Fischer-Tropsch catalystmay be in the range of from about 0.005 and 0.050, in the range of fromabout 0.0075 and 0.0125, or about 0.010. Copper may serve as aninduction promoter. In preferred embodiments, the weight ratio of Cu:Feis about 1:100.

The catalyst may be an iron Fischer-Tropsch catalyst comprisingstructural promoter. The structural promoter may significantly reducethe breakdown of the catalyst in a SBCR (slurry bubble column reactor).The structural promoter may comprise silica, and may enhance thestructural integrity during activation and operation of the catalyst. Inembodiments, the catalyst comprises a mass ratio of SiO₂:Fe of less thanabout 1:100 when the structural promoter comprises silica and less thanabout 8:100 when the structural promoter comprises silica sol.

In embodiments, the at least one structural promoter is selected fromoxides of metals and metalloids and combinations thereof. The structuralpromoter may be referred to as a binder, a support material, or astructural support.

Depending on the level of structural promoter comprising silicate andthe preselected alpha, i.e. the polymerization probability desired, theweight ratio of K:Fe may be from about 0.5:100 to about 6.5:100, fromabout 0.5:100 to about 2:100, or about 1:100.

In embodiments wherein the structural promoter comprises silica sol, theweight ratio of iron to potassium is in the range of from about 100:1 toabout 100:5. In embodiments, the weight ratio of iron to potassium is inthe range of from about 100:2 to about 100:6. In embodiments, the weightratio of iron to potassium is in the range of from about 100:3 to about100:5. In embodiments, the weight ratio of iron to potassium is in therange of from about 100:4 to about 100:5. In some preferred embodiments,the weight ratio of iron to potassium is in the range of from about100:2 to about 100:4. In embodiments, the weight ratio of iron topotassium about 100:3. In embodiments, the weight ratio of iron topotassium is about 100:5.

In embodiments wherein the structural promoter comprises silica sol, theweight ratio of iron to copper may be in the range of from about 100:1to about 100:7. In some embodiments, the weight ratio of iron to copperis in the range of from about 100:1 to about 100:5. More preferably, theweight ratio of iron to copper is in the range of from about 100:2 toabout 100:6. Still more preferably, the weight ratio of iron to copperis in the range of from about 100:3 to about 100:5. In some preferredembodiments, the weight ratio of iron to copper is in the range of fromabout 100:2 to about 100:4. In other specific embodiments, the weightratio of iron to copper is about 100:5. In yet other specificembodiments, the weight ratio of iron to copper is about 100:3.

Broadly, in embodiments, wherein the structural promoter is silica sol,the iron to SiO₂ weight ratio may be in the range of from about 100:1 toabout 100:8; alternatively, in the range of from 100:1 to 100:7. Inembodiments, wherein the structural promoter is silica, the iron to SiO₂weight ratio may be in the range of from about 100:2 to about 100:6. Inembodiments, the weight ratio of iron to silica is in the range of fromabout 100:3 to about 100:5. In embodiments, wherein the structuralpromoter is silica, the iron to SiO₂ weight ratio is about 100:5. Inembodiments, wherein the structural promoter is silica, the iron to SiO₂weight ratio may be in the range of from about 100:3 to about 100:7;alternatively, in the range of from about 100:4 to about 100:6. Inembodiments, the catalyst comprises an Fe:Cu:K:SiO₂ mass ratio of about100:4:3:5.

In embodiments, the FT catalyst is a cobalt-based catalyst. Inembodiments, the catalyst comprises cobalt, and optionally a co-catalystand/or promoter, supported on a support wherein the cobalt loading is atleast or about 5, 10, 15, 20, 25, 28, 30, 32, 35, or 40 percent byweight. In embodiments, the cobalt loading is in the range of from about5 to about 50% by weight, from about 10 to about 50% by weight, fromabout 15 to about 50% by weight, from about 20 to about 50% by weight,from about 25 to about 50% by weight, from about 28 to about 50% byweight, from about 30 to about 50% by weight, or from about 32 to about50% by weight. The metal dispersion for the catalytically active metal(e.g., Co, and optionally co-catalyst and/or promoter) of the catalystmay be in the range of from about 1 to about 30%, from about 2 to about20%, or from about 3 to about 20%. In embodiments, the co-catalyst isselected from the group consisting of Fe, Ni, Ru, Re, Os, oxidesthereof, and mixtures of two or more thereof. In embodiments, thecatalyst comprises at least one promoter selected from the groupconsisting of Group IA, IIA, IIIB or IVB metals, oxides thereof,lanthanide metals and oxides thereof, and actinide metals and oxidesthereof. In embodiments, the promoter is selected from the groupconsisting of Li, B, Na, K, Rb, Cs, Mg, Ca, Sr, Ba, Sc, Y, La, Ac, Ti,Zr, La, Ac, Ce, Th, oxides thereof, and mixtures of two or more thereof.The co-catalyst may be employed at a concentration in the range of fromabout 0 to about 10% by weight based on the total weight of the catalyst(i.e., the weight of catalyst, co-catalyst, promoter and support) orfrom about 0.1 to about 5% by weight. The promoter may be employed at aconcentration of up to about 10% by weight based on the total weight ofthe catalyst, and in one embodiment about 0.1 to about 5% by weight.

In embodiments, the catalyst comprises cobalt supported by alumina; theloading of cobalt being at least about 25% by weight, at least about 28%by weight, at least about 30% by weight, or at least about 32% byweight; and the cobalt dispersion is at least about 3%, at least about5%, or at least about 7%.

In embodiments, the catalyst used in the disclosed method is a FTcatalyst as described in and/or formed via the multiple impregnationstep method described in U.S. Pat. No. 7,084,180, the disclosure ofwhich is hereby incorporated herein in its entirety for all purposes notcontrary to this disclosure. The catalyst can comprises at least onecatalytic metal (i.e., Co, Fe) at a loading level of about 20% by weightor more, about 25% by weight or more, about 28% by weight or more, about30% by weight or more, about 32% by weight or more, about 35% by weightor more, about 37% by weight or more, or about 40% by weight or more.

In embodiments, the FT catalyst utilized is an iron-based catalyst andsubjecting the synthesis gas to reaction conditions whereby liquidhydrocarbons are produced comprises contacting the synthesis gas withcatalyst at a temperature in the range of from about 200° C. (392° F.)to about 300° C. (572° F.), from about 220° C. (428° F.) to about 275°C. (527° F.) or from about 240° C. (464° F.) to about 260° C. (500° F.).In embodiments, the temperature of the FT synthesis is a temperature ofgreater than or equal to about 200° C. (392° F.), 220° C. (428° F.) or250° C. (482° F.). In embodiments, the FT synthesis is carried out at apressure in the range of from about 100 psig (689.5 kPa) to about 1000psig (6894.8 kPa), from about 200 psig (1379 kPa) to about 500 psig(3447.4 kPa) or from about 300 psig (2068.4 kPa) to about 400 psig(2757.9 kPa). In embodiments, the FT synthesis is carried out at apressure of greater than or equal to about 100 psig (689.5 kPa), about300 psig (2068.4 kPa), or about 350 psig (2413.2 kPa). In embodiments,the FT synthesis is carried by introducing the synthesis gas into an FTproduction apparatus 700, as described hereinabove. The FT productionapparatus comprises an FT synthesis reactor. In embodiments, the FTsynthesis reactor is a slurry bubble column reactor. In embodiments, theresidence time in the FT synthesis reactor is in the range of from about1 s to about 3000 s, from about 10 s to about 500 s or from about 100 sto about 300 s. In embodiments, the FT synthesis is carried out for aresidence time of about 100 s, about 200 s, or about 300 s.

Converting synthesis gas into product can further comprise separatingtailgas and/or spent catalyst/product from the conversion product at1020. In embodiments, converting synthesis gas into product furthercomprises separating tailgas from the primary product(s). For example, atailgas comprising methane, hydrogen, carbon monoxide and/or carbondioxide may be separated from the synthesis gas conversion product (e.g.liquid hydrocarbons). In embodiments, converting synthesis gas intoproduct further comprises separating spent catalyst/wax from the liquidhydrocarbons. During FT synthesis of liquid hydrocarbons, spent catalystand associated wax is routinely removed from the slurry process. Suchcatalyst can be separated from the primary liquid product via anysuitable methods known in the art, for example via centrifugation,filtration, magnetic separation, or a combination thereof. Suchseparated spent catalyst and any product that is separated therewith isreferred to herein as spent catalyst/product and may be recycled, asdiscussed further hereinbelow, in step 1100.

Recycling at Least One Component 1100.

In embodiments, the disclosed method of producing synthesis gasconversion product according to this disclosure comprises recycling atleast one component from converting at 1000 for reuse in producingadditional synthesis gas 800.

In embodiments, recycling at least one component 1100 comprisesrecycling at least a portion of the tailgas produced while convertingthe synthesis gas to product at 1000. The tailgas produced duringconversion of synthesis gas to product can be recycled for use as fuelin the reforming step. In this manner, the ‘waste’ tailgas can beutilized to benefit within the system. As discussed hereinabove, one ormore tailgas recycle lines 770 may fluidly connect the FT synthesisapparatus 700 with one or more burners 404 of reformer 400. As discussedhereinabove, carbon dioxide may be removed from the tailgas prior torecycle for use as fluid. This may be particularly desirable as itenables sequestration of additional carbon dioxide within thebiorefinery. All or a portion of the tailgas exiting synthesis gasconversion apparatus 700 via tailgas outlet line 760 can be introducedinto a carbon dioxide removal apparatus, described hereinabove, and/orrecycled to reformer 400.

In embodiments, recycling at least one component 1100 comprisesrecycling at least a portion of the separated spent catalyst/product andsubjecting it to reforming conditions along with carbonaceous feedmaterial to produce additional synthesis gas from the product (e.g.liquid hydrocarbons) therein. For example, one or more catalyst/productrecycle lines 755 may be used to introduce spent catalyst/product intothe reformer, for example via carbonaceous feed line 250 and/or mixingapparatus 300. In this manner, additional value can be obtained from thereformable material separated with the spent catalyst, thus improvingthe overall liquid yields of the biorefinery. Additionally, in thismanner, disposal of material is reduced and the spent catalyst(substantially free of associated product) can be separated (e.g. withthe ash) for disposal. At least a portion of the spent catalyst/productseparated from the primary conversion product(s) at 1020 can thus berecycled through the coiled tubes of the reformer for production ofadditional synthesis gas therefrom.

While the preferred embodiments of the invention have been shown anddescribed, modifications thereof can be made by one skilled in the artwithout departing from the spirit and teachings of the invention. Theembodiments described and the examples provided herein are exemplaryonly, and are not intended to be limiting. Many variations andmodifications of the invention disclosed herein are possible and arewithin the scope of the invention. Accordingly, the scope of protectionis not limited by the description set out above, but is only limited bythe claims which follow, that scope including all equivalents of thesubject matter of the claims.

The discussion of a reference is not an admission that it is prior artto the present invention, especially any reference that may have apublication date after the priority date of this application. Thedisclosures of all patents, patent applications, and publications citedherein are hereby incorporated herein by reference in their entirety, tothe extent that they provide exemplary, procedural, or other detailssupplementary to those set forth herein.

What is claimed is:
 1. A system for the production of conversionproducts from synthesis gas, the system comprising: a mixing apparatusconfigured for mixing steam with at least one carbonaceous material toproduce a reformer feedstock, wherein the mixing apparatus comprises oneor more cylindrical vessel having a conical bottom section, an inlet forsuperheated steam within the conical bottom section and an inlet for theat least one carbonaceous material at or near the top of the cylindricalvessel, wherein the one or more cylindrical vessel is a pressure vesselconfigured for operation at a pressure in the range of from about 5 psig(34.5 kPa) to about 50 psig (344.7 kPa); a reformer configured toproduce, from the reformer feedstock, a reformer product comprisingsynthesis gas comprising hydrogen and carbon monoxide, and alsoproducing a hot flue gas; a synthesis gas conversion apparatusconfigured to catalytically convert at least a portion of the synthesisgas in the reformer product into synthesis gas conversion productcomprising primarily liquid hydrocarbons, and to separate, from thesynthesis gas conversion product, a spent catalyst stream comprisingspent catalyst and associated synthesis gas conversion product separatedtherewith, and a tailgas comprising at least one gas selected from thegroup consisting of carbon monoxide, carbon dioxide, hydrogen andmethane; and a recycle line fluidly connecting the synthesis gasconversion apparatus with the mixing apparatus, whereby at least aportion of the spent catalyst stream can be introduced into the mixingapparatus.
 2. The system of claim 1 further comprising a recycle linefluidly connecting the synthesis gas conversion apparatus with at leastone burner of the reformer, whereby at least a portion of the tailgascan be combusted to provide heat.
 3. The system of claim 1 furthercomprising: a steam superheater configured to produce superheated steamvia heat transfer from the hot flue gas to a steam superheater feedsteam, and comprising an outlet for a reduced-temperature flue gas; andone or more steam generator configured to produce the steam superheaterfeed steam via heat transfer from the reformer product gas and thereduced-temperature flue gas.
 4. The system of claim 3 wherein the oneor more steam generator configured to produce the steam superheater feedsteam via heat transfer from the reformer product gas and thereduced-temperature flue gas comprises a reformer flue gas and reformereffluent steam generator comprising an inlet for the reduced-temperatureflue gas, an inlet for the reformer product gas, and an inlet for water,and configured to produce the steam superheater feed steam via heattransfer to the water from the reformer product gas and thereduced-temperature flue gas.
 5. The system of claim 3 wherein the oneor more steam generator configured to produce the steam superheater feedsteam via heat transfer from the reformer product gas and thereduced-temperature flue gas comprises a reformer effluent steamgenerator comprising an inlet for the reformer product gas and an inletfor water, and configured to produce a reformer effluent steam generatorproduct via heat transfer from the reformer product gas to the water;and a reformer flue gas steam generator comprising an inlet for thereformer effluent steam generator product and the reduced-temperatureflue gas, and configured to produce the steam superheater feed steam viaheat transfer to the reformer effluent steam generator product from thereduced-temperature flue gas.
 6. The system of claim 1 wherein the atleast one carbonaceous material comprises biomass.
 7. The system ofclaim 1 wherein the pressure vessel is operable at a pressure of about 5psig (34.5 kPa) to 45 psig (310.3 kPa).
 8. The system of claim 1 whereinthe one or more cylindrical vessel is configured for operation at apressure in the range of from about 30 psig (206.8 kPa) to about 50 psig(344.7 kPa).
 9. The system of claim 1 wherein the metallurgy of thereformer allows operation at a reformer temperature greater than orequal to about 1700° F. (926° C.) and a reformer pressure greater thanor equal to about 5 psig (34.5 kPa).
 10. The system of claim 1 whereinthe reformer comprises: a cylindrical vessel containing a plurality ofvertically-oriented coiled tubes fluidly connected with the mixingapparatus such that reformer feedstock may be introduced thereto; atleast one burner configured to combust fuel to provide heat for thereforming and produce the hot flue gas; at least one outlet for reformerproduct; and at least one outlet for the hot flue gas.
 11. The system ofclaim 10 wherein each of the plurality of vertically-oriented coiledtubes has a vertical height in the range of from about 40 feet (12.2 m)to about 100 feet (30.5 m) and a coil length at least 4 times thevertical height.
 12. The system of claim 10 wherein the total coillength is in the range of from about 4 to about 25 times the verticalheight.
 13. The system of claim 12 wherein the total coil length is inthe range of from about 4 to about 12 times the vertical height.
 14. Thesystem of claim 10 wherein at least a portion of the plurality ofvertically-oriented coiled tubes have an inside diameter (ID) in therange of from about 2 inches (5.1 cm) to about 4 inches (10.2 cm). 15.The system of claim 10 wherein the metallurgy of the coiled tubes allowsoperation at a reformer pressure greater than or equal to about 45 psig(310.3 kPa).
 16. The system of claim 10 wherein the at least one burneris positioned substantially at, near, or below the bottom of thecylindrical vessel.
 17. The system of claim 10 wherein the reformercomprises from about 1 to about 10 burners.
 18. The system of claim 10wherein outlets of each of the vertically-oriented coiled tubes aremanifolded into an outlet for the reformer product, wherein the manifoldis positioned at, near, or below the bottom of the cylindrical vessel.19. The system of claim 10 wherein the at least one outlet for flue gasis positioned at the top of the cylindrical vessel.
 20. The system ofclaim 1 further comprising feed preparation apparatus configured tocomminute the at least one carbonaceous material, to dry the at leastone carbonaceous material, or both.
 21. The system of claim 20 whereinthe feed preparation apparatus comprises at least one grinder and atleast one separator configured to provide a carbonaceous material havingan average particle diameter of less than about 3/16^(th) of an inch(0.47 cm).
 22. The system of claim 1 wherein the synthesis gasconversion apparatus is configured to catalytically convert at least aportion of the synthesis gas in the reformer product into synthesis gasconversion product via contact of the at least a portion of thesynthesis gas with an iron-based Fischer-Tropsch catalyst.
 23. A systemfor the production of conversion products from synthesis gas, the systemcomprising: a mixing apparatus configured for mixing steam with at leastone carbonaceous material to produce a reformer feedstock, wherein themixing apparatus comprises a pressure vessel configured for operation ata pressure of greater than about 30 psig (206.8 kPa); a reformerconfigured to produce, from the reformer feedstock, a reformer productcomprising synthesis gas comprising hydrogen and carbon monoxide, andalso producing a hot flue gas; a synthesis gas conversion apparatusconfigured to catalytically convert at least a portion of the synthesisgas in the reformer product into synthesis gas conversion productcomprising primarily liquid hydrocarbons, and to separate, from thesynthesis gas conversion product, a spent catalyst stream comprisingspent catalyst and associated synthesis gas conversion product separatedtherewith, and a tailgas comprising at least one gas selected from thegroup consisting of carbon monoxide, carbon dioxide, hydrogen andmethane; a recycle line fluidly connecting the synthesis gas conversionapparatus with the mixing apparatus, whereby at least a portion of thespent catalyst stream can be introduced into the mixing apparatus; asteam superheater configured to produce superheated steam via heattransfer from the hot flue gas to a steam superheater feed steam, andcomprising an outlet for a reduced-temperature flue gas; and one or moresteam generator configured to produce the steam superheater feed steamvia heat transfer from the reformer product gas and thereduced-temperature flue gas.
 24. The system of claim 23 wherein thesteam superheater feed steam comprises saturated steam.